Addition of chlorine to regenerator

ABSTRACT

A process for economically converting carbo-metallic oils to lighter products. The carbo-metallic oils contain 650° F. and material which is characterized by a carbon residue on pyrolysis of at least about 1 and a Nickel Equivalents of heavy metals content of at least about 4 parts per million. This process comprises flowing the carbo-metallic oil together with particulate cracking catalyst through a progressive flow type reactor having an elongated reaction chamber, which is at least in part vertical or inclined, for a predetermined vapor riser residence time in the range of about 0.5 to about 10 seconds, at a temperature of about 900° to about 1400° F., and under a pressure of about 10 to about 50 pounds per square inch absolute sufficient for causing a conversion per pass in the range of about 50% to about 90% while producing coke in amounts in the range of about 6 to about 14% by weight based on fresh feed, and laying down coke on the catalyst in amounts in the range of about 0.3 to about 3% by weight. The spent, coke-laden catalyst from the stream of hydrocarbons formed by vaporized feed and resultant cracking products is separated and regenerated in one or more regeneration beds in one or more regeneration zones by burning the coke on the spent catalyst with oxygen in the presence of chlorine. The bed density is in the range of about 25 to about 50 pounds per cubic foot and the bed or beds are sustained by fluidization gas containing combustion-supporting gas having a linear velocity in the range of about 0.2 to about 4 feed per second. The chlorine reduces the CO 2  /CO ratio, thereby reducing the heat output of the regenerator. The catalyst particles are retained in the regeneration zone or zones in contact with the combustion-supporting gas for an average total residence time in said zone or zones of about 5 to about 30 minutes to reduce the level of carbon on the catalyst to about 0.25% by weight or less. The regenerated catalyst is recycled to the reactor and contacted with fresh carbo-metallic oil.

This application is a continuation-in-part application of copendingapplications Ser. Nos. 94,092 (now U.S. Pat. No. 4,332,673) and 94,216(now U.S. Pat. No. 4,341,624), each filed Nov. 14, 1979 and eachentitled "Carbo-Metallic Oil Conversion."

DESCRIPTION

1. Technical Field

This invention relates to processes for converting heavy hydrocarbonoils into ligher fractions, and especially to processes for convertingheavy hydrocarbons containing high concentrations of coke precursors andheavy metals into gasoline and other liquid hydrocarbon fuels.

2. Background Art

In general, gasoline and other liquid hydrocarbon fuels boil in therange of about 100° to about 650° F. However, the crude oil from whichthese fuels are made contains a diverse mixture of hydrocarbons andother compounds which vary widely in molecular weight and therefore boilover a wide range. For example, crude oils are known in which 30 to 60%or more of the total volume of oil is composed of compounds boiling attemperatures above 650° F. Among these are crudes in which about 10% toabout 30% or more of the total volume consists of compounds so heavy inmolecular weight that they boil above 1025° F. or at least will not boilbelow 1025° F. at atmospheric pressure.

Because these relatively abundant high boiling components of crude oilare unsuitable for inclusion in gasoline and other liquid hydrocarbonfuels, the petroleum refining industry has developed processes forcracking or breaking the molecules of the high molecular weight, highboiling compounds into smaller molecules which do boil over anappropriate boiling range. The cracking process which is most widelyused for this purpose is known as fluid catalytic cracking (FCC).Although the FCC process has reached a highly advanced state, and manymodified forms and variations have been developed, their unifying factoris that a vaporized hydrocarbon feedstock is caused to crack at anelevated temperature in contact with a cracking catalyst that issuspended in the feedstock vapors. Upon attainment of the desired degreeof molecular weight and boiling point reduction the catalyst isseparated from the desired products.

Crude oil in the natural state contains a variety of materials whichtend to have quite troublesome effects on FCC processes, and only aportion of these troublesome materials can be economically removed fromthe crude oil. Among these troublesome materials are coke precursors(such as asphaltenes, polynuclear aromatics, etc.), heavy metals (suchas nickel, vanadium, iron, copper, etc.), lighter metals (such assodium, potassium, etc.), sulfur, nitrogen and others. Certain of these,such as the lighter metals, can be economically removed by desaltingoperations, which are part of the normal procedure for pretreating crudeoil for fluid catalytic cracking. Other materials, such as cokeprecursors, asphaltenes and the like, tend to break down into cokeduring the cracking operation, which coke deposits on the catalyst,impairing contact between the hydrocarbon feedstock and the catalyst,and generally reducing its potency or activity level. The heavy metalstransfer almost quantitatively from the feedstock to the catalystsurface.

If the catalyst is reused again and again for processing additionalfeedstock, which is usually the case, the heavy metals can accumulate onthe catalyst to the point that they unfavorably alter the composition ofthe catalyst and/or the nature of its effect upon the feedstock. Forexample, vanadium tends to form fluxes with certain components ofcommonly used FCC catalysts, lowering the melting point of portions ofthe catalyst particles sufficiently so that they begin to sinter andbecome ineffective cracking catalysts. Accumulations of vanadium andother heavy metals, especially nickel, also "poison" the catalyst. Theytend in varying degrees to promote excessive dehydrogenation andaromatic condensation, resulting in excessive production of carbon andgases with consequent impairment of liquid fuel yield. An oil such as acrude or crude fraction or other oil that is particularly abundant innickel and/or other metals exhibiting similar behavior, while containingrelatively large quantities of coke precursors, is referred to herein asa carbo-metallic oil, and represents a particular challenge to thepetroleum refiner.

In general the coke-forming tendency or coke precursor content of an oilcan be ascertained by determining the weight percent of carbon remainingafter a sample of that oil has been pyrolyzed. The industry accepts thisvalue as a measure of the extent to which a given oil tends to formnon-catalytic coke when employed as feedstock in a catalytic cracker.Two established tests are recognized, the Conradson Carbon andRamsbottom Carbon tests, the former being described in ASTM D189-76 andthe latter being described in ASTM Test No. D524-76. In conventional FCCpractice, Conradson carbon values on the order of about 0.05 to about1.0 are regarded as indicative of acceptable feed. The present inventionis concerned with the use of hydrocarbon feedstocks which have higherConradson carbon values and thus exhibit substantially greater potentialfor coke formation than the usual feeds.

Since the various heavy metals are not of equal catalyst poisoningactivity, it is convenient to express the poisoning activity of an oilcontaining a given poisoning metal or metals in terms of the amount of asingle metal which is estimated to have equivalent poisoning activity.

Thus, the heavy metals content of an oil can be expressed by thefollowing formula (patterned after that of W. L. Nelson in Oil and GasJournal, page 143, Oct. 23, 1961) in which the content of each metalpresent is expressed in parts per million of such metal, as metal, on aweight basis, based on the weight of feed: ##EQU1##

According to conventional FCC practice, the heavy metal content offeedstock for FCC processing is controlled at a relatively low level,e.g. about 0.25 ppm Nickel Equivalents or less. The present invention isconcerned with the processing of feedstocks containing metalssubstantially in excess of this value and which therefore have asignificantly greater potential for accumulating on and posoningcatalyst.

The above formula can also be employed as a measure of the accumulationof heavy metals on cracking catalyst, except that the quantity of metalemployed in the formula is based on the weight of catalyst (moisturefree basis) instead of the weight of feed. In conventional FCC practice,in which a circulating inventory of catalyst is used again and again inthe processing of fresh feed, with periodic or continuing minor additionand withdrawal of fresh and spent catalyst, the metal content of thecatalyst is maintained at a level which may for example be in the rangeof about 200 to about 600 ppm Nickel Equivalents. The process of thepresent invention is concerned with the use of catalyst having asubstantially larger metals content, and which therefore has a muchgreater than normal tendency to promote dehydrogenation, aromaticcondensation, gas production or coke formation. Therefore, such highermetals accumulation is normally regarded as quite undesirable in FCCprocessing.

There has been a long standing interest in the conversion ofcarbo-metallic oils into gasoline and other liquid fuels. For example,in the 1950s it was suggested that a variety of carbo-metallic oilscould be successfully converted to gasoline and other products in theHoudresid process. Turning from the FCC mode of operation, the Houdresidprocess employed catalyst particles of "granular size" (much larger thanconventional FCC catalyst particle size) in a compact gravitating bed,rather than suspending catalyst particles in feed and product vapors ina fluidized bed.

Although the Houdresid process obviously represented a step forward indealing with the effects of metal contamination and coke formation oncatalyst performance, is productivity was limited. Because its operationwas uneconomical, the first Houdresid unit is no longer operating. Thus,for the 25 years which have passed since the Houdresid process was firstintroduced commercially, the art has continued its arduous search forsuitable modifications or alternatives to the FCC process which wouldpermit commercially successful operation on reduced crude and the like.During this period a number of proposals have been made; some have beenused commercially to a certain extent.

Several proposals involve treating the heavy oil feed to remove themetal therefrom prior to cracking, such as by hydrotreating, solventextraction and complexing with Friedel-Crafts catalysts, but thesetechniques have been criticized as unjustified economically. Anotherproposal employs a combination cracking process having "dirty oil" and"clean oil" units. Still another proposal blends residual oil with gasoil and controls the quantity of residual oil in the mixture in relationto the equilibrium flash vaporization temperature at the bottom of theriser type cracker unit employed in the process. Still another proposalsubjects the feed to a mild preliminary hydrocracking or hydrotreatingoperation before it is introduced into the cracking unit. It has alsobeen suggested to contact a carbo-metallic oil such as reduced crudewith hot taconite pellets to produce gasoline. This is a small samplingof the many proposals which have appeared in the patent literature andtechnical reports.

Notwithstanding the great effort which has been expended and the factthat each of these proposals overcomes some of the difficultiesinvolved, conventional FCC practice today bears mute testimony to thedearth of carbo-metallic oil-cracking techniques that are botheconomical and highly practical in terms of technical feasibility. Somecrude oils are relatively free of coke precursors or heavy metals orboth, and the troublesome components of crude oils are for the most partconcentrated in the highest boiling fractions. Accordingly, it has beenpossible to largely avoid the problems of coke precursors and heavymetals by sacrificing the liquid fuel yield which would be potentiallyavailable from the highest boiling fractions. More particularly,conventional FCC practice has employed as feedstock that fraction ofcrude oil which boils at about 650° F. to about 1000° F., such fractionbeing relatively free of coke precursors and heavy metal contamination.Such feedstock, known as "vacuum gas oil" (VG0) is generally preparedfrom crude oil by distilling off the fractions boiling below about 650°F. at atmospheric pressure and then separating by further vacuumdistillation from the heavier fractions a cut boiling between about 650°F. and about 900° to 1025° F.

The vacuum gas oil is used as feedstock for conventional FCC processing.The heavier fractions are normally employed for a variety of otherpurposes, such as for instance production of asphalt, residual fuel oil,#6 fuel oil, or marine Bunker C fuel oil, which represents a great wasteof the potential value of this portion of the crude oil, especially inlight of the great effort and expense which the art has been willing toexpend in the attempt to produce generally similar materials from coaland shale oils. The present invention is aimed at the simultaneouscracking of these heavier fractions containing substantial quantities ofboth coke precursors and heavy metals, and possibly other troublesomecomponents, in conjunction with the lighter oils, thereby increasing theoverall yield of gasoline and other hydrocarbon liquid fuels from agiven quantity of crude. As indicated above, the present invention by nomeans constitutes the first attempt to develop such a process, but thelong standing recognition of the desirability of cracking carbo-metallicfeedstocks, along with the slow progress of the industry toward doingso, shows the continuing need for such a process. It is believed thatthe present process is uniquely advantageous for dealing with theproblem of treating such carbo-metallic oils in an economically andtechnically sound manner.

One method for cracking these high boiling fractions, named ReducedCrude Conversion (RCC) after a particularly common and usefulcarbo-metallic feed, is disclosed in copending application Ser No.94,216 filed Nov. 14, 1979 (now U.S. Pat. No. 4,341,624) for"Carbo-Metallic Oil Conversion". The oils disclosed as capable of beingcracked by the method of that application are carbo-metallic oils atleast about 70 percent of which boil above 650° F. and containing acarbon residue on pyrolysis of at least about 1 and at least about 4parts per million of nickel equivalents of heavy metals. Examples ofthese oils are crude oils, topped crudes, reduced crudes, residua, andextracts from solvent de-asphalting.

The cracking reaction for the method disclosed in application Ser. No.94,216 is sufficiently severe to convert 50% or more of the feedstock togasoline per pass and produce coke in the amount of 6 to 14% by weightbased on weight of fresh feed. In a typical RCC cracking process theratio of weight of catalyst to weight of feedstock is from about 3 toabout 18, coke is laid down on the catalyst in amounts in the range ofabout 0.3 to about 3 percent by weight based on the weight of thecatalyst, and heavy metals accumulate on the catalyst to a concentrationof from about 3000 to about 30,000 ppm nickel equivalents.

The unusually large amount of coke which deposits on the catalyst incarbo-metallic oil processing presents critical problems, the primaryproblem arising from the fact that the reactions in the regeneratorwhich convert coke to water, carbon monoxide and carbon dioxide arehighly exothermic. Using a carbo-metallic feed with its unusually highcontent of coke precursors as compared to FCC feeds, can increase theamount of coke burned in the regenerator and the temperature in theregenerator to the point that regeneration temperatures become excessiveif there is thorough burning of coke. Excessive temperatures canpermanently deactivate the catalyst and/or damage the regeneratingequipment.

The heat of combustion of coke depends upon the concentration ofhydrogen in the coke and the ratio of CO₂ to CO in the products ofcombustion. Carbon produces 13,910 BTU per pound when burned to CO₂ andonly 3,962 BTU per pound when burned to CO. Hydrogen produces 61,485 BTUper pound when burned to H₂ O. The heats of combustion of coke for theserepresentative levels of hydrogen and four different ratios of CO₂ /COare given in the following table:

                  TABLE I                                                         ______________________________________                                        Heat of Combustion BTU/lb Coke                                                           Percent Hydrogen                                                   CO.sub.2 /CO Ratio                                                                         2           4       6                                            ______________________________________                                        0.5           8,362       9,472  10,582                                       1.0                      11,038  12,083                                       3.0                              14,446                                       4.0          12,912              14,894                                       ______________________________________                                    

These problems encountered in regenerating catalysts coated with a highconcentration of coke may be exacerbated when catalysts of the zeoliteor molecular sieve type are used. These catalysts, which are crystallinealuminosilicates made up of tetra-coordinated aluminum atoms associatedthrough oxygen atoms with silicon atoms in the crystalline structure,tend to be less susceptible than prior amorphous silica-aluminacatalysts in respect to loss of cracking activity upon extended exposureto high temperatures. Also, they have been shown to be more adverselyaffected by coke in respect to loss of cracking activity, than arecertain other catalysts, such as for example the non-zeolite,silica-alumina catalysts.

Various methods have been used to control the temperature in theregeneration zone including cooling by heat exchangers external to theregenerator (see U.S. Pat. No. 2,394,710), cooling by injecting steam orwater into an upper, dilute phase zone of a regenerator (see U.S. Pat.No. 3,909,392), and controlling the oxidation reaction by controllingthe amount of oxygen present (see U.S. Pat. No. 3,161,583). These andother methods which have been proposed control the temperature of theregenerator for conventional FCC feedstocks having Conradson carbonresidues below about one percent. However, processes for convertingfeedstocks containing Conradson carbon residues greater than about twopercent require a method of heat control other than those normally used.

U.S. applications Ser. Nos. 94,092 (now U.S. Pat. No. 4,332,673) and94,227, (now U.S. Pat. No. 4,354,923), filed Nov. 14, 1979, discloseprocesses for the conversion of carbo-metallic oils to liquid fuel inwhich various regeneration techniques are employed that assist incontrolling the heat load in the regeneration step. It is thought thatthe ratio of CO₂ to CO may be decreased to no more than about 4 andpreferably to less than about 3 in order to reduce the amount of energyreleased within the regenerator, while optionally providing a flue gashigh enough in CO content to be a useful fuel. However, it is difficultto increase this ratio to a sufficiently high level and simultaneouslybring the coke level on the catalyst to a desirably low level i.e., lessthan about 0.1 percent, especially with RCC feeds since the heavy metalsi.e., vanadium, iron, nickel and copper, which accumulate on thecatalyst with the coke act as oxidation promoters to convert CO to CO₂.Levels of heavy metals greater than about 3000 ppm and especiallygreater than about 10,000 ppm present severe difficulties in reducingthe CO₂ /CO ratio. This emphasizes the need for oxidation suppressorswhich do not reduce the amount of coke burned. As will be appreciatedthe carbo-metallic oils can vary widely in their Conradson carboncontent. Such varying content of carbon residue in the feedstock, alongwith variations in riser operating conditions such as catalyst to oilratio and others, can result in wide variation of the percent coke foundon the spent catalyst. Accordingly, where the feed and riser operatingconditions are such as to produce rather large coke yields,necessitating the burning of very substantial amounts of coke from thecatalyst in regeneration, such as at least about 0.5 weight percentbased on the catalyst, or more, additional measures for controlling theheat load in the regenerator may prove useful. One purpose of thepresent invention is to meet this need.

SUMMARY OF THE INVENTION

Accordingly it is one object of this invention to provide a catalyticcracking method for converting carbo-metallic oils to liquid fuelswherein the heat generated within the catalyst regenerator is reduced.

It is another object to provide a carbo-metallic oil conversion processinvolving controlling the heat evolved within the regenerator whereinthe weight of coke burned is at least about 0.5% of the weight of thecatalyst.

It is still another object to provide a carbo-metallic oil conversionprocess wherein the ratio of CO/CO₂ in the regenerator flue gases isincreased while effecting substantially complete combustion of coke.

In accordance with this invention a process is provided for convertingcarbo-metallic feedstocks containing 650° F. material and having aresidue on pyrolysis of at least about one and containing at least aboutfour ppm nickel equivalents of heavy metals comprising bringing saidfeedstock in contact with a cracking catalyst in a progressive flowreactor under conditions whereby light products and coke are formed andcoke and heavy metals deposit onto the catalyst. The spent catalyst isregenerated by contacting it with an oxygen-containing gas in thepresence of chlorine under conditions whereby at least a portion of saidcoke is converted to carbon oxides, and the regenerated catalyst isrecycled to the reactor for contact with fresh feed.

The addition of chlorine to the gases in the regenerator results in anincrease in the CO/CO₂ ratio and a concomitant decrease in the heatproduced in the regenerator. While this process is useful for feedstockshaving a Conradson carbon value of about one, it is particularly usefulfor processing feedstocks having Conradson carbon values of at leastabout two, and it is especially useful for processing feedstocks havingConradson carbon contents greater than about 6. The invention isespecially useful in meeting the heat removal requirements involved inregenerating catalyst which has been used in cracking carbo-metallicoils containing high concentrations of coke precursors and heavy metalsand which are therefore heavily loaded with coke and heavy metals.

Depending on how the process of the invention is practised, one or moreof the following additional advantages may be realized. If desired andpreferably the process may be operated without added hydrogen in thereaction chamber. If desired, and preferably, the process may beoperated without prior hydrotreating of the feed and/or without otherprocesses of removal of asphaltenes or metals from the feed, and this istrue even where the carbo-metallic oil as a whole contains more thanabout 4, or more than about 5 or even more than about 5.5 ppm NickelEquivalents by weight of heavy metal and has a carbon residue onpyrolysis greater than about 1% or greater than about 2% by weight.Moreover, all of the converter feed, as above described, may be crackedin one and the same conversion chamber. The cracking reaction may becarried out with a catalyst which has previously been used (recycled,except for such replacement as required to compensate for normal lossesand deactivation) to crack a carbo-metallic feed under the abovedescribed conditions. Heavy hydrocarbons not cracked to gasoline in afirst pass may be recycled with or without hydrotreating for furthercracking in contact with the same kind of feed in which they were firstsubjected to cracking conditions, and under the same kind of conditions;but operation in a substantially once-through or single pass mode (e.g.less than about 15% by volume of recycle based on volume of fresh feed)is preferred.

BRIEF DESCRIPTION OF THE DRAWING

The FIGURE is a schematic design of catalyst regeneration apparatus andassociated cracking apparatus which may be used in carrying out thisinvention.

BEST AND VARIOUS OTHER ILLUSTRATIVE MODES FOR CARRYING OUT THE INVENTION

In carrying out this invention the chlorine is added to the system sothat it is carried into an oxidation zone of the regenerator. It may bemixed with theoxidizing gas added to the regenerator or it may be addedspearately. In the preferred method of carrying out this invention thechlorine is mixed with the oxidizing gas.

While the mechanism by which chlorine increases the CO/CO₂ ratio is notknown, it is hypothesized that the chlorine limits the oxidation of COto CO₂. Under typical conditions within a catalyst regenerator whereinthe temperature is in the range of about 1050° F. the followingreactions may take place:

    C+O.sub.2 →CO.sub.2                                 (1)

    2C+O.sub.2 →2CO                                     (2)

    2CO+O.sub.2 →2CO.sub.2                              (3)

    CO.sub.2 +C→2CO                                     (4)

All these reactions occur to some extent in a catalyst regenerator, andwhen oxygen is present in a stoichiometric amount or greater the productwill be predominantly CO₂. Reaction (3) is probably initiated andpropagated by free radicals, the formation of which is believed to beretarded by chlorine.

The chlorine is removed from the system with the flue gases andtherefore must be continuously replenished at a rate sufficiently highto provide a concentration useful in retarding the oxidation of CO toCO₂. Chlorine in concentrations in the regenerator gases of at leastabout 100 ppm is effective in increasing the CO/CO₂ ratio, whilesignificantly lower concentrations, as for example 50 ppm, do not appearto change this ratio. An increase in chlorine concentration increasesthe change in the CO/CO₂ ratio. However, factors such as cost andenvironmental considerations establish a preferred upper limit to thechlorine concentration at about 400 ppm. The concentration of chlorinein the regenerator is preferably in the range from about 100 to about400 ppm and most preferably is in the range of about 100 to about 300ppm based on the total weight of gases in the regenerator.

The present invention is notable in providing a simple, relativelystraightforward and highly productive approach to the conversion ofcarbo-metallic feed such as reduced crude or the like to various lighterproducts such as gasoline. The carbo-metallic feed comprises or iscomposed of oil which boils above about 650° F. Such oil, or at leastthe 650° F.+ portion thereof, is characterized by a heavy metal contentof at least about 4, preferably more than about 5, and most preferablyat least about 5.5 ppm of Nickel Equivalents by weight and by a carbonresidue on pyrolysis of at least about 1% and more preferably at leastabout 2% by weight. In accordance with the invention, the carbo-metallicfeed, in the form of a pumpable liquid, is brought into contact with hotconversion catalyst in a weight ratio of catalyst to feed in the rangeof about 3 to about 18 and preferably more than about 6.

The feed in said mixture undergoes a conversion step which includescracking while the mixture of feed and catalyst is flowing through aprogressive flow type reactor. The feed, catalyst, and other materialsmay be introduced at one or more points. The reactor includes anelongated reaction chamber which is at least partly vertical or inclinedand in which the feed material, resultant products and catalyst aremaintained in contact with one another while flowing as a dilute phaseor stream for a predetermined riser residence time in the range of about0.5 to about 10 seconds.

The reaction is conducted at a temperature of about 900° to about 1400°F., measured at the reaction chamber exit, under a total pressure ofabout 10 to about 50 psia (pounds per square inch absolute) underconditions sufficiently severe to provide a conversion per pass in therange of about 50% or more and to lay down coke on the catalyst in anamount in the range of about 0.3 to about 3% by weight and preferably atleast about 0.5%. The overall rate of coke production, based on weightof fresh feed, is in the range of about 4 to about 14% by weight.

At the end of the predetermined residence time, the catalyst isseparated from the products, is stripped to remove high boilingcomponents and other entrained or adsorbed hydrocarbons and is thenregenerated with oxygen-containing combustion-supporting gas underconditions of time, temperature and atmosphere sufficient to reduce thecarbon on the regenerated catalyst to about 0.25% or less and preferablyabout 0.05% or less by weight.

Depending on how the process of the invention is practiced, one or moreof the following additional advantages may be realized. If desired, andpreferably, the process may be operated without added hydrogen in thereaction chamber. If desired, and preferably, the process may beoperated without prior hydrotreating of the feed and/or without otherprocess of removal of asphaltenes of metals from the feed, and this istrue even where the carbo-metallic oil as a whole contains more thanabout 4, or more than about 5 or even more than about 5.5 ppm NickelEquivalents by weight of heavy metal and has a carbon residue onpyrolysis greater than about 1%, greater than about 1.4% or greater thanabout 2% by weight. Moreover, all of the converter feed, as abovedescribed, may be cracked in one and the same conversion chamber. Thecracking reaction may be carried out with a catalyst which haspreviously been used (recycled, except for such replacement as requiredto compensate for normal losses and deactivation) to crack acarbo-metallic feed under the above described conditions. Heavyhydrocarbons not cracked to gasoline in a first pass may be recycledwith or without hydrotreating for further cracking in contact with thesame kind of feed in which they were first subjected to crackingconditions, and under the same kind of conditions; but operation in asubstantially once-through or single pass mode (e.g. less than about 15%by volume of recycle based on volume of fresh feed) is preferred.

According to one preferred embodiment or aspect of the invention, at theend of the predetermined residence time referred to above, the catalystis projected in a direction established by the elongated reactionchamber or an extension thereof, while the products, having lessermomentum, are caused to make an abrupt change of direction, resulting inan abrupt, substantially instantaneous ballistic separation of productsfrom catalyst. The thus separated catalyst is then stripped, regeneratedand recycled to the reactor as above described.

According to another preferred embodiment or aspect of the invention,the converter feed contains 650° F.+ material which has not beenhydrotreated and is characterized in part by containing at least about5.5 parts per million of nickel equivalents of heavy metals. Theconverter feed is brought together not only with the above mentionedcracking catalyst, but also with additional gaseous material includingsteam whereby the resultant suspension of catalyst and feed alsoincludes gaseous material wherein the ratio of the partial pressure ofthe added gaseous material relative to that of the feed is in the rangeof about 0.25 to about 4.0. The vapor residence time is in the range ofabout 0.5 to about 3 seconds when practicing this embodiment or aspectof the invention. This preferred embodiment or aspect and the onereferred to in the preceeding paragraph may be used in combination withone another or separately.

According to another preferred embodiment or aspect of the invention,the carbo-metallic feed is not only brought into contact with thecatalyst, but also with one or more additional materials includingparticularly liquid water in a weight ratio relative to feed rangingfrom about 0.04 to about 0.25, more preferably about 0.04 to about 0.2and still more preferably about 0.05 to about 0.15. Such additionalmaterials, including the liquid water, may be brought into admixturewith the feed prior to, during or after mixing the feed with theaforementioned catalyst, and either after or, preferably, before,vaporization of the feed. The feed, catalyst and water (e.g. in the formof liquid water or in the form of steam produced by vaporization ofliquid water in contact with the feed) are introduced into theprogressive flow type reactor, which may or may not be a reactorembodying the above described ballistic separation, at one or morepoints along the reactor. While the mixture of feed, catalyst and steamproduced by vaporization of the liquid water flows through the reactor,the feed undergoes the above mentioned conversion step which includescracking. The feed material, catalyst, steam and resultant products aremaintained in contact with one another in the above mentioned elongatedreaction chamber while flowing as a dilute phase or stream for the abovementioned predetermined riser residence time which is in the range ofabout 0.5 to about 10 seconds.

The present invention provides a process for the continuous catalyticconversion of a wide variety of carbo-metallic oils to lower molecularweight products, while maximizing production of highly valuable liquidproducts, and making it possible, if desired, to avoid vacuumdistillation and other expensive treatments such as hydrotreating. Theterm "oils," includes not only those predominantly hydrocarboncompositions which are liquid at room temperature (i.e., 68° F.), butalso those predominantly hydrocarbon compositions which are asphalts ortars at ambient temperature but liquify when heated to temperatures inthe range of up to about 800° F. The invention is applicable tocarbo-metallic oils, whether of petroleum origin or not. For example,provided they have the requisite boiling range, carbon residue onpyrolysis and heavy metals content, the invention may be applied to theprocessing of such widely diverse materials as heavy bottoms from crudeoil, heavy bitumen crude oil, those crude oils known as "heavy crude"which approximate the properties of reduced crude, shale oil, tar sandextract, products from coal liquification and solvated coal, atmosphericand vacuum reduced crude, extracts and/or bottoms (raffinate) fromsolvent de-asphalting, aromatic extract from lube oil refining, tarbottoms, heavy cycle oil, slop oil, other refinery waste streams andmixtures of the foregoing. Such mixtures can for instance be prepared bymixing available hydrocarbon fractions, including oils, tars, pitchesand the like. Also, powdered coal may be suspended in the carbo-metalliccoil. Persons skilled in the art are aware of techniques fordemetalation of carbo-metallic oils, and demetalated oils may beconverted using the invention; but it is an advantage of the inventionthat it can employ as feedstock carbo-metallic oils that have had noprior demetalation treatment. Likewise, the invention can be applied tohydrotreated feedstocks; but it is an advantage of the invention that itcan successfully convert carbo-metallic oils which have hadsubstantially no prior hydrotreatment. However, the preferredapplication of the process is to reduced crude, i.e., that fraction ofcrude oil boiling at and above 650° F., alone or in admixture withvirgin gas oils. While the use of material that has been subjected toprior vacuum distillation is not excluded, it is an advantage of theinvention that it can satisfactorily process material which has had noprior vacuum distillation, thus saving on capital investment andoperating costs as compared to conventional FCC processes that require avacuum distillation unit.

In accordance with the invention one provides a carbo-metallic oilfeedstock, at least about 70%, more preferably at least about 85% andstill more preferably about 100% (by volume) of which boils at and aboveabout 650° F. All boiling temperatures herein are based on standardatmospheric pressure conditions. In carbo-metallic oil partly or whollycomposed of material which boils at and above about 650° F., suchmaterial is referred to herein as 650° F.+ material; and 650° F.+material which is part of or has been separated from an oil containingcomponents boiling above and below 650° F. may be referred to as a 650°F.+ fraction. but the terms "boils above" and "650° F.+" are notintended to imply that all of the material characterized by said termswill have the capability of boiling. The carbo-metallic oilscontemplated by the invention may contain material which may not boilunder any conditions; for example, certain asphalts and asphaltenes maycrack thermally during distillation, apparently without boiling. Thus,for example, when it is said that the feed comprises at least about 70%by volume of material which boils above about 650° F., it should beunderstood that the 70% in question may include some material which willnot boil or volatilize at any temperature. These non-boilable materialswhen present, may frequently or for the most part be concentrated inportions of the feed which do not boil below about 1000° F., 1025° F. orhigher. Thus, when it is said that at least about 10%, more preferablyabout 15% and still more preferably at least about 20% (by volume) ofthe 650° F.+ fraction will not boil below about 1000° F. or 1025° F., itshould be understood that all or any part of the material not boilingbelow about 1000° or 1025° F., may or may not be volatile at and abovethe indicated temperatures.

Preferably, the contemplated feeds, or at least the 650° F.+ materialtherein, have a carbon residue on pyrolysis of at least about 2 orgreater. For example, the Conradson carbon content may be in the rangeof about 2 to about 12 and most frequently at least about 4. Aparticularly common range is about 4 to about 8. Those feeds having aConradson carbon content greater than about 6 especially need specialmeans for controlling excess heat.

Preferably, the feed has an average composition characterized by anatomic hydrogen to carbon ratio in the range of about 1.2 to about 1.9,and preferably about 1.3 to about 1.8.

The carbo-metallic feeds employed in accordance with the invention, orat least the 650° F.+ material therein, may contain at least about 4parts per million of Nickel Equivalents, as defined above, of which atleast about 2 parts per million is nickel (as metal, by weight).Carbo-metallic oils within the above range can be prepared from mixturesof two or more oils, some of which do and some of which do not containthe quantities of Nickel Equivalents and nickel set forth above. Itshould also be noted that the above values of Nickel Equivalents andnickel represent time-weighted averages for a substantial period ofoperation of the conversion unit, such as one month, for example. Itshould also be noted that the heavy metals have in certain circumstancesexhibited some lessening of poisoning tendency after repeated oxidationsand reductions on the catalyst, and the literature describes criteriafor establishing "effective metal" values. For example, see the articleby Cimbalo, et al, entitled "Deposited Metals Poison FCC Catalyst," Oiland Gas Journal, May 15, 1972, pp 112-122, the contents of which areincorporated herein by reference. If considered necessary or desirable,the contents of Nickel Equivalents and nickel in the carbo-metallic oilsprocessed according to the invention may be expressed in terms of"effective metal" values. Notwithstanding the gradual reduction inpoisoning activity noted by Cimbalo, et al, the regeneration of catalystunder normal FCC regeneration conditions may not, and usually does not,severely impair the dehydrogenation, demethanation and aromaticcondensation activity of heavy metals accumulated on cracking catalyst.

It is known that about 0.2 to about 5 weight percent of "sulfur" in theform of elemental sulfur and/or its compounds (but reported as elementalsulfur based on the weight of feed) appears in FCC feeds and that thesulfur and modified forms of sulfur can find their way into theresultant gasoline product and, where lead is added, tend to reduce itssusceptibility to octane enhancement. Sulfur in the product gasolineoften requires sweetening when processing high sulfur containing crudes.To the extent that sulfur is present in the coke, it also represents apotential air pollutant since the regenerator burns it to SO₂ and SO₃.However, we have found that in our process the sulfur in the feed is onthe other hand able to inhibit heavy metal activity by maintainingmetals such as Ni, V, Cu and Fe in the sulfide form in the reactor.These sulfides are much less active than the metals themselves inpromoting dehydrogenation and coking reactions. Accordingly, it isacceptable to carry out the invention with a carbo-metallic oil havingat least about 0.3%, acceptably more than about 0.8% and more acceptablyat least about 1.5% by weight of sulfur in the 650° F.+ fraction.

The carbo-metallic oils useful in the invention may and usually docontain significant quantities of compounds containing nitrogen, asubstantial portion of which may be basic nitrogen. For example, thetotal nitrogen content of the carbo-metallic oils may be a least about0.05% by weight. Since cracking catalysts owe their cracking activity toacid sites on the catalyst surface or in its pores, basicnitrogen-containing compounds may temporarily neutralize these sites,poisoning the catalyst. However, the catalyst is not permanently damagedsince the nitrogen can be burned off the catalyst during regeneration,as a result of which the acidity of the active sites is restored.

The carbo-metallic oils may also include significant quantities ofpentane insolubles, for example at least about 0.5% by weight, and moretypically 2% or more or even about 4% or more. These may include forinstance asphaltenes and other materials.

Alkali and alkaline earth metals generally do not tend to vaporize inlarge quantities under the distillation conditions employed indistilling crude oil to prepare the vacuum gas oils normally used as FCCfeedstocks. Rather, these metals remain for the most part in the"bottoms" fraction (the non-vaporized high boiling portion) which mayfor instance be used in the production of asphalt or other by-products.However, reduced crude and other carbo-metallic oils are in many casesbottoms products, and therefore may contain significant quantities ofalkali and alkaline earth metals such as sodium. These metals depositupon the catalyst during cracking. Depending on the composition of thecatalyst and magnitude of the regeneration temperatures to which it isexposed, these metals may undergo interactions and reactions with thecatalyst (including the catalyst support) which are not normallyexperienced in processing VGO under conventional FCC processingconditions. If the catalyst characteristics and regeneration conditionsso require, one will of course take the necessary precautions to limitthe amounts of alkali and alkaline earth metal in the feed, which metalsmay enter the feed not only as brine associated with the crude oil inits natural state, but also as components of water or steam which aresupplied to the cracking unit. Thus, careful desalting of the crude usedto prepare the carbo-metallic feed may be important when the catalyst isparticularly susceptible to alkali and alkaline earth metals. In suchcircumstances, the content of such metals (hereinafter collectivelyreferred to as "sodium") in the feed can be maintained at about 1 ppm orless, based on the weight of the feedstock. Alternatively, the sodiumlevel of the feed may be keyed to that of the catalyst, so as tomaintain the sodium level of the catalyst which is in use substantiallythe same as or less than that of the replacement catalyst which ischarged to the unit.

According to a particularly preferred embodiment of the invention, thecarbo-metallic oil feedstock constitutes at least about 70% by volume ofmaterial which boils above about 650° F., and at least about 10% of thematerial which boils above about 650° F. will not boil below about 1025°F. The average composition of this 650° F.+ material may be furthercharacterized by: (a) an atomic hydrogen to carbon ratio in the range ofabout 1.3 to about 1.8: (b) a Conradson carbon value of at least about2; (c) at least about four parts per million of Nickel Equivalents, asdefined above, of which at least about two parts per million is nickel(as metal, by weight); and (d) at least one of the following: (i) atleast about 0.3% by weight of sulfur, (ii) at least about 0.05% byweight of nitrogen, and (iii) at least about 0.5% by weight of pentaneinsolubles. Very commonly, the preferred feed will include all of (i),(ii), and (iii), and other components found in oils of petroleum andnon-petroleum origin may also be present in varying quantities providingthey do not prevent operation of the process.

Although there is no intention of excluding the possiblity of using afeedstock which has previously been subjected to some cracking, thepresent invention has the definite advantage that it can successfullyproduce large conversions and very substantial yields of liquidhydrocarbon fuels from carbo-metallic oils which have not been subjectedto any substantial amount of cracking. Thus, for example, andpreferably, at least about 85%, more preferably at least about 90% andmost preferably substantially all of of the carbo-metallic feedintroduced into the present process is oil which has not previously beencontacted with cracking catalyst under cracking conditions. Moreover,the process of the invention is suitable for operation in asubstantially once-through or single pass mode. Thus, the volume ofrecycle, if any, based on the volume of fresh feed is preferably about15% or less and more preferably about 10% or less.

In general, the weight ratio of catalyst to fresh feed (feed which hasnot previously been exposed to cracking catalyst under crackingconditions) used in the process is in the range of about 3 to about 18.Preferred and more preferred ratios are about 4 to about 12, morepreferably about 5 to about 10 and still more preferably about 6 toabout 10, a ratio of about 10 presently being considered most nearlyoptimum. Within the limitations of product quality requirements,controlling the catalyst to oil ratio at relatively low levels withinthe aforesaid ranges tends to reduce the coke yield of the process,based on fresh feed.

In conventional FCC processing of VGO, the ratio between the number ofbarrels per day of plant through-put and the total number of tons ofcatalyst undergoing circulation throughout all phases of the process canvary widely. For purposes of this disclosure, daily plant through-put isdefined as the number of barrels of fresh feed boiling above about 650°F. which that plant processes per average day of operation to liquidproducts boiling below about 430° F. For example, in one commerciallysuccessful type of FCC-VGO operation, about 8 to about 12 tons ofcatalyst are under circulation in the process per 1000 barrels per dayof plant through-put. In another commercially successful process, thisratio is in the range of about 2 to 3. While the present invention maybe practiced in the range of about 2 to about 30 and more typicallyabout 2 to about 12 tons of catalyst inventory per 1000 barrels of dailyplant through-put, it is preferred to carry out the process of thepresent invention with a very small ratio of catalyst weight to dailyplant through-put. More specifically, it is preferred to carry out theprocess of the present invention with an inventory of catalyst that issufficient to contact the feed for the desired residence time in theabove indicated catalyst to oil ratio while minimizing the amount ofcatalyst inventory, relative to plant through-put, which is undergoingcirculation or being held for treatment in other phases of the processsuch as, for example, stripping, regeneration and the like. Thus, moreparticularly, it is preferred to carry out the process of the presentinvention with about 2 to about 5 and more preferably about 2 tons ofcatalyst inventory or less per thousand barrels of daily plantthrough-put.

In the practice of the invention, catalyst may be added continuously orperiodically, such as, for example, to make up for normal losses ofcatalyst from the system. Moreover, catalyst addition may be conductedin conjunction with withdrawal of catalyst, such as, for example, tomaintain or increase the average activity level of the catalyst in theunit. For example, the rate at which virgin catalyst is added to theunit may be in the range of about 0.1 to about 3, more preferably about0.15 to about 2, and most preferably about 0.2 to about 1.5 pounds perbarrel of feed. If on the other hand equilibrium catalyst from FCCoperation is to be utilized, replacement rates as high as about 5 poundsper barrel can be practiced. Where circumstances are such that thecatalyst employed in the unit is below average in resistance todeactivation and/or conditions prevailing in the unit are such as topromote more rapid deactivation, one may employ rates of additiongreater than those stated above; but in the opposite circumstances,lower rates of addition may be employed. By way of illustration, if aunit were operated with a metal(s) loading of 5000 ppm Ni+V in parts byweight on equilibrium catalyst, one might for example employ areplacement rate of about 2.7 pounds of catalyst introduced for eachbarrel (42 gallons) of feed processed. However, operation at a higherlevel such as 10,000 ppm Ni+V on catalyst would enable one tosubstantially reduce the replacement rate, such as for example to about1.3 pounds of catalyst per barrel of feed. Thus, the levels of metal(s)on catalyst and catalyst replacement rates may in general berespectively increased and decreased to any value consistent with thecatalyst activity which is available and desired for conducting theprocess.

Without wishing to be bound by any theory, it appears that a number offeatures of the process to be described in greater detail below, suchas, for instance, the residence time and optional mixing of steam withthe feedstock, tend to restrict the extent to which cracking conditionsproduce metals in the reduced state on the catalyst from heavy metalsulfide(s), sulfate(s) or oxide(s) deposited on the catalyst particlesby prior exposures to carbo-metallic feedstocks and regenerationconditions. Thus, the process appears to afford significant control overthe poisoning effect of heavy metals on the catalyst even when theaccumulations of such metals are quite substantial.

Accordingly, the process may be practised with catalyst bearingaccumulations of heavy metal(s) in the form of elemental metal(s),oxide(s), sulfide(s) or other compounds which heretofore would have beenconsidered quite intolerable in conventional FCC-VGO operations. Thus,operation of the process with catalyst bearing heavy metalsaccumulations in the range of about 3,000 or more ppm NickelEquivalents, on the average, is contemplated. The concentration ofNickel Equivalents of metals on catalyst can range up to about 50,000ppm or higher. More specifically, the accumulation may be in the rangeof about 3,000 to about 30,000 ppm, preferably in the range of 3,000 to20,000 ppm, and more particularly about 3,000 to about 12,000 ppm.Within these ranges just mentioned, operation at metals levels of about4,000 or more, about 5,000 or more, or about 7,000 or more ppm can tendto reduce the rate of catalyst replacement required. The foregoingranges are based on parts per million of Nickel Equivalents, in whichthe metals are expressed as metal, by weight, measured on and based onregenerated equilibrium catalyst. However, in the event that catalyst ofadequate activity is available at very low cost, making feasible veryhigh rates of catalyst replacement, the carbo-metallic oil could beconverted to lower boiling liquid products with catalyst bearing lessthan 3,000 ppm Nickel Equivalents of heavy metals. For example, onemight employ equilibrium catalyst from another unit, for example, an FCCunit which has been used in the cracking of a feed, e.g. vacuum gas oil,having a carbon residue on pyrolysis of less than 1 and containing lessthan about 4 ppm Nickel Equivalents of heavy metals.

In any event, the equilibrium concentration of heavy metals in thecirculating inventory of catalyst can be controlled (includingmaintained or varied as desired or needed) by manipulation of the rateof catalyst addition discussed above. Thus, for example, addition ofcatalyst may be maintained at a rate which will control the heavy metalsaccumulation on the catalyst in one of the ranges set forth above.

In general, it is preferred to employ a catalyst having a relativelyhigh level of cracking activity, providing high levels of conversion andproductivity at low residence times. The conversion capabilities of thecatalyst may be expressed in terms of the conversion produced duringactual operation of the process and/or in terms of conversion producedin standard catalyst activity tests. For example, it is preferred toemploy catalyst which, in the course of extended operation underprevailing process conditions, is sufficiently active for sustaining alevel of conversion of at least about 50% and more preferably at leastabout 60%. In this connection, conversion is expressed in liquid volumepercent, based on fresh feed. Also, for example, the preferred catalystmay be defined as one which, in its virgin or equilibrium state,exhibits a specified activity expressed as a percentage in terms of MAT(micro-activity test) conversion. For purposes of the present inventionthe foregoing percentage is the volume percentage of standard feedstockwhich a catalyst under evaluation will convert to 430° F. end pointgasoline, lighter products and coke at 900° F., 16 WHSV (weight hourlyspace velocity, calculated on a moisture free basis, using cleancatalyst which has been dried at 1100° F., weighed and then conditioned,for a period of at least 8 hours at about 25° C. and 50% relativehumidity, until about one hour or less prior to contacting the feed) and3 C/O (catalyst to oil weight ratio) by ASTM D-32 MAT test D-3907-80,using an appropriate standard feedstock, e.g. a sweet light primary gasoil, such as that used by Davison, Division of W. R. Grace, having thefollowing analysis and properties:

    ______________________________________                                        API Gravity at 60° F., degrees                                                             31.0                                                      Specific Gravity at 60° F., g/cc                                                           0.8708                                                    Ramsbottom Carbon, wt. %                                                                          0.09                                                      Conradson Carbon, wt. % (est.)                                                                    0.04                                                      Carbon, wt. %       84.92                                                     Hydrogen, wt. %     12.94                                                     Sulfur, wt. %       0.68                                                      Nitrogen, ppm       305                                                       Viscosity at 100° F., centistokes                                                          10.36                                                     Watson K Factor     11.93                                                     Aniline Point       182                                                       Bromine No.         2.2                                                       Paraffins, Vol. %   31.7                                                      Olefins, Vol. %     1.6                                                       Naphthenes, Vol. %  44.0                                                      Aromatics, Vol. %   22.7                                                      Average Molecular Weight                                                                          284                                                       Nickel              Trace                                                     Vanadium            Trace                                                     Iron                Trace                                                     Sodium              Trace                                                     Chlorides           Trace                                                     B S & W             Trace                                                     ______________________________________                                        Distillation        ASTM D-1160                                               ______________________________________                                        IBP                 445                                                       10%                 601                                                       30%                 664                                                       50%                 701                                                       70%                 734                                                       90%                 787                                                       FBP                 834                                                       ______________________________________                                    

The gasoline end point and boiling temperature-volume percentrelationships of the product produced in the MAT conversion test may forexample be determined by simulated distillation techniques, for example,modifications of gas chromatograph "Sim-D," ASTM D-2887-73. The resultsof such simulations are in reasonable agreement with the resultsobtained by subjecting larger samples of material to standard laboratorydistillation techniques. Conversion is calculated by subtracting from100 the volume percent (based on fresh feed) of those products heavierthan gasoline which remain in the recovered product.

On page 935-937 of Hougen and Watson, Chemical Process Principles, JohnWiley & Sons, Inc., N.Y. (1947), the concent of "Activity Factors" isdiscussed. This concept leads to the use of "relative activity" tocompare the effectiveness of an operating catalyst against a standardcatalyst, as developed by Shankland and Schmitkons. Relative activitymeasurements facilitate recognition of how the quantity requirements ofvarious catalysts differ from one another. Thus, relative activity is aratio obtained by dividing the weight of a standard or referencecatalyst which is or would be required to produce a given level ofconversion, as compared to the weight of an operating catalyst (whetherproposed or actually used) which is or would be required to produce thesame level of conversion in the same or equivalent feedstock under thesame or equivalent conditions. Said ratio of catalyst weights may beexpressed as a numerical ratio, but preferably is converted to apercentage basis. The standard catalyst is preferably chosen from amongcatalysts useful for conducting the present invention, such as forexample zeolite fluid cracking catalysts, and is chosen for its abilityto produce a predetermined level of conversion in a standard feed underthe conditions of temperature, WHSV, catalyst to oil ratio and otherconditions set forth in the preceding description of the MAT conversiontest and in ASTM D-32 MAT test D-3907-80. Conversion is the volumepercentage of feedstock that is converted to 430° F. endpoint gasoline,lighter products and coke. For standard feed, one may employ theabove-mentioned light primary gas oil, or equivalent.

For purposes of conducting relative activity determinations, one mayprepare a "standard catalyst curve," a chart or graph of conversion (asabove defined) vs. reciprocal WHSV for the standard catalyst andfeedstock. A sufficient number of runs is made under ASTM D-3907-80conditions (as modified above) using standard feedstock at varyinglevels of WHSV to prepare an accurate "curve" of conversion vs. WHSV forthe standard feedstock. This curve should traverse all or substantiallyall of the various levels of conversion including the range ofconversion within which it is expected that the operating catalyst willbe tested. From this curve, one may establish a standard WHSV for testcomparisons and standard value of reciprocal WHSV corresponding to thatlevel of conversion which has been chosen to represent 100% relativeactivity in the standard catalyst. For purposes of the presentdisclosure the aforementioned reciprocal WHSV and level of conversionare, respectively, 0.0625 and 75%. In testing an operating catalyst ofunknown relative activity, one conducts a sufficient number of runs withthat catalyst under D-3907-80 conditions (as modified above) toestablish the level of conversion which is or would be produced with theoperating catalyst at standard reciprocal WHSV. Then, using theabove-mentioned standard catalyst curve, one establishes a hypotheticalreciprocal WHSV constituting the reciprocal WHSV which would have beenrequired, using the standard catalyst, to obtain the same level ofconversion which was or would be exhibited, by the operating catalyst atstandard WHSV. The relative activity may then be calculated by dividingthe hypothetical reciprocal WHSV by the reciprocal standard WHSV, whichis 1/16, or 0.0625. The result is relative activity expressed in termsof a decimal fraction, which may then be multiplied by 100 to convert topercent relative activity. In applying the results of thisdetermination, a relative activity of 0.5, or 50%, means that it wouldtake twice the amount of the operating catalyst to give the sameconversion as the standard catalyst, i.e., the production catalyst is50% as active as the reference catalyst.

The catalyst may be introduced into the process in its virgin form or,as previously indicated, in other than virgin form; e.g. one may useequilibrium catalyst withdrawn from another unit, such as catalyst thathas been employed in the cracking of a different feed. Whethercharacterized on the basis of MAT conversion activity or relativeactivity, the preferred catalysts may be described on the basis of theiractivity "as introduced" into the process of the present invention, oron the basis of their "as withdrawn" or equilibrium activity in theprocess of the present invention, or on both of these bases. A preferredactivity level of virgin and non-virgin catalyst "as introduced" intothe process of the present invention is at least about 60% by MATconversion, and preferably at least about 20%, more preferably at leastabout 40% and still more preferably at least about 60% in terms ofrelative activity. However, it will be appreciated that, particularly inthe case of non-virgin catalysts supplied at high addition rates, loweractivity levels may be acceptable. An acceptable "as withdrawn" orequilibrium activity level of catalyst which has been used in theprocess of the present invention is at least about 20% or more, butabout 40% or more and preferably about 60% or more are preferred valueson a relative activity basis, and an activity level of 60% or more on aMAT conversion basis is also contemplated. More preferably, it isdesired to employ a catalyst which will, under the conditions of use inthe unit, establish an equilibrium activity at or above the indicatedlevel. The catalyst activities are determined with catalyst having lessthan 0.01 coke, e.g. regenerated catalyst.

One may employ any hydrocarbon cracking catalyst having the aboveindicated conversion capabilities. A particularly preferred class ofcatalysts includes those which have pore structures into which moleculesof feed material may enter for adsorption and/or for contact with activecatalytic sites within or adjacent the pores. Various types of catalystsare available within this classification, including for example thelayered silicates, e.g. smectites. Although the most widely availablecatalysts within this classification are the well-knownzeolite-containing catalysts, non-zeolite catalysts are alsocontemplated.

The preferred zeolite-containing catalysts may include any zeolite,whether natural, semi-synthetic or synthetic, alone or in admixture withother materials which do not significantly impair the suitability of thecatalyst, provided the resultant catalyst has the activity and porestructure referred to above. For example, if the virgin catalyst is amixture, it may include the zeolite component associated with ordispersed in a porous refractory inorganic oxide carrier, in such casethe catalyst may for example contain about 1% to about 60%, morepreferably about 15% to about 50%, and most typically about 20 to about45% by weight, based on the total weight of catalyst (water free basis)of the zeolite, the balance of the catalyst being the porous refractoryinorganic oxide alone or in combination with any of the known adjuvantsfor promoting or suppressing various desired and undesired reactions.For a general explanation of the genus of zeolite, molecular sievecatalysts useful in the invention, attention is drawn to the disclosuresof the articles entitled "Refinery Catalysts Are a Fluid Business" and∓Making Cat Crackers Work On Varied Diet," appearing respectively in theJuly 26, 1978 and Sept. 13, 1978 issues of Chemical Week magazine. Thedescriptions in the aforementioned publications are incorporated hereinby reference.

For the most part, the zeolite components of the zeolite-containingcatalysts, will be those which are known to be useful in FCC crackingprocesses. In general, these are crystalline aluminosilicates, typicallymade up of tetra coordinated aluminum atoms associated through oxygenatoms with adjacent silicon atoms in the crystal structure. However, theterm "zeolite" as used in this disclosure contemplates not onlyaluminosilicates, but also substances in which the aluminum has beenpartly or wholly replaced, such as for instance by gallium and/or othermetal atoms, and further includes substances in which all or part of thesilicon has been replaced, such as for instance by germanium. Titaniumand zirconium substitution may also be practiced.

Most zeolites are prepared or occur naturally in the sodium form, sothat sodium cations are associated with the electronegative sites in thecrystal structure. The sodium cations tend to make zeolites inactive andmuch less stable when exposed to hydrocarbon conversion conditions,particularly high temperatures. Accordingly, the zeolite may be ionexchanged, and where the zeolite is a component of a catalystcomposition, such ion exchanging may occur before or after incorporationof the zeolite as a component of the composition. Suitable cations forreplacement of sodium in the zeolite crystal structure include ammonium(decomposable to hydrogen), hydrogen, rare earth metals, alkaline earthmetals, etc. Various suitable ion exchange procedures and cations whichmay be exchanged into the zeolite crystal structures are well known tothose skilled in the art.

Examples of the naturally occuring crystalline aluminosilicate zeoliteswhich may be used as or included in the catalyst for the presentinvention are faujasite, mordenite, clinoptilote, chabazite, analcite,crionite, as well as levynite, dachiardite, paulingite, noselite,ferriorite, heulandite, scolccite, stibite, harmotome, phillipsite,brewsterite, flarite, datolite, gmelinite, caumnite, leucite, lazurite,scaplite, mesolite, ptolite, nephline, matrolite, offretite andsodalite.

Examples of the synthetic crystalline aluminosilicate zeolites which areuseful as or in the catalyst for carrying out the present invention areZeolite X, U.S. Pat. No. 2,882,244, Zeolite Y, U.S. Pat. No. 3,130,007;and Zeolite A, U.S. Pat. No. 2,882,243; as well as Zeolite B, U.S. Pat.No. 3,008,803; Zeoite D, Canada Pat. No. 661,981; Zeolite E, Canada Pat.No. 614,495; Zeolite F, U.S. Pat. No. 2,996,358; Zeolite H. U.S. Pat.No. 3,010,789; Zeolite J., U.S. Pat. No. 3.011,869; Zeolite L, BelgianPat. No. 575,177; Zeolite M., U.S. Pat. No. 2,995,423, Zeolite O, U.S.Pat. No. 3,140,252; Zeolite Q, U.S. Pat. No. 2,991,151; Zeolite S, U.S.Pat. No. 3,054,657, Zeolite T, U.S. Pat. No. 2,950,952; Zeolite W, U.S.Pat. No. 3,012,853; Zeolite Z, Canada Pat. No. 614,495; and ZeoliteOmega, Canada Pat. No. 817,915. Also, ZK-4HJ, alpha beta and ZSM-typezeolites are useful. Moreover, the zeolites described in U.S. Pat. Nos.3,140,249, 3,140,253, 3,944,482 and 4,137,151 are also useful, thedisclosures of said patents being incorporated herein by reference.

The crystalline aluminosilicate zeolites having a faujasite-type crystalstructure are particularly preferred for use in the present invention.This includes particularly natural faujasite and Zeolite X and ZeoliteY.

The crystalline aluminosilicate zeolites, such as synthetic faujasite,will under normal conditions crystallize as regularly shaped, discreteparticles of about one to about ten microns in size, and, accordingly,this is the size range frequently found in commercial catalysts whichcan be used in the invention. Preferably, the particle size of thezeolites is from about 0.1to about 10 microns and more preferably isfrom about 0.1 to about 2 microns or less. For example, zeolitesprepared in situ from calcined kaoline may be characterized by evensmaller crystallites. Crystalline zeolites exhibit both an interior andan exterior surface area, which we have defined as "portal" surfacearea, with the largest portion of the total surface area being internal.By portal surface area, we refer to the outer surface of the zeolitecrystal through which reactants are considered to pass in order toconvert to lower boiling products. Blockages of the internal channelsby, for example, coke formation, blockages of entrance to the internalchannels by deposition of coke in the portal surface area, andcontamination by metals poisoning, will greatly reduce the total zeolitesurface area. Therefore, to minimize the effect of contamination andpore blockage, crystals larger than the normal size cited above arepreferably not used in the catalysts of this invention.

Commercial zeolite-containing catalysts are available with carrierscontaining a variety of metal oxides and combination thereof, includingfor example silica, alumina, magnesia, and mixtures thereof and mixturesof such oxides with clays as e.g. described in U.S. Pat. No. 3,034,948.One may for example select any of the zeolite-containing molecular sievefluid cracking catalysts which are suitable for producting of gasolinefrom vacuum gas oils. However, certain advantages may be attained byjudicious selection of catalysts having marked resistance to metals. Ametal resistant zeolite catalyst is, for instance, described in U.S.Pat. No. 3,944,482, in which the catalyst contains 1-40 weight percentof a rare earth-exchanged zeolite, the balance being a refractory metaloxide having specified pore volume and size distribution. Othercatalysts described as "metals-tolerant" are described in the abovementioned Cimbalo et al article.

In general, it is preferred to employ catalysts having an over-allparticle size in the range of about 5 to about 160, more preferablyabout 40 to about 120, and most preferably about 40 to about 80 microns.For example, a useful catalyst may have a skeletal density of about 150pounds per cubic foot and an average particle size of about 60-70microns, with less than 10% of the particles having a size less thanabout 40 microns and less than 80% having a size less than about 50-60microns.

Although a wide variety of other catalysts, including bothzeolite-containing and non-zeolite-containing may be employed in thepractice of the invention the following are examples of commerciallyavailable catalysts which may be employed in practicing the invention:

                  TABLE II                                                        ______________________________________                                        Spe-                                                                          cific      Weight Percent                                                     Sur-       zeolite                                                            face       Con-                                                               m.sup.2 /g tent    Al.sub.2 O.sub.3                                                                      SiO.sub.2                                                                          Na.sub.2 O                                                                          Fe.sub.2 O                                                                          TiO.sub.2                         ______________________________________                                        AGZ-290 300    11.0    29.5  59.0 0.40  0.11  0.59                            GRZ-1   162    14.0    23.4  69.0 0.10  0.4   0.9                             CCZ-220 129    11.0    34.6  60.0 0.60  0.57  1.9                             Super DX                                                                              155    13.0    31.0  65.0 0.80  0.57  1.6                             F-87    240    10.0    44.0  50.0 0.80  0.70  1.6                             FOX-90  240    8.0     44.0  52.0 0.65  0.65  1.1                             HFZ 20  310    20.0    59.0  40.0 0.47  0.54  2.75                            HEZ 55  210    19.0    59.0  35.2 0.60  0.60  2.5                             ______________________________________                                    

The AGZ-290, GRZ-1, CCZ-220 and Super DX catalysts referred to above areproducts of W. R. Grace and Co. F-87 and FOX-90 are products of Filtrol,while HFZ-20 and HEZ-55 are products of Engelhard/Houdry. The above areproperties of virgin catalyst and, except in the case of zeolitecontent, are adjusted to a water free basis, i.e. based on materialignited at 1750° F. The zeolite content is derived by comparison of theX-ray intensities of a catalyst sample and of a standard materialcomposed of high purity sodium Y-zeolite in accordance with draft #6,Jan. 9, 1978, of proposed ASTM Standard Method entitled "Determinationof the Faujasite Content of a Catalyst."

Among the above mentioned commercially available catalysts, the Super Dfamily and especially a catalyst designated GRZ-1 are particularlypreferred. For example, Super DX has given particularly good resultswith Arabian light crude. The GRZ-1, although substantially moreexpensive than the Super DX at present, appears somewhat moremetals-tolerant.

Although not yet commercially available, it is believed that the bestcatalysts for carrying out the present invention will be those which,according to proposals advanced by Dr. William P. Hettinger, Jr. and Dr.James E. Lewis, are characterized by matrices with feeder pores havinglarge minimum diameters and large mouths to facilitate diffusion of highmolecular weight molecules through the matrix to the portal surface areaof molecular sieve particles within the matrix. Such matrices preferablyalso have a relatively large pore volume in order to soak up unvaporizedportions of the carbo-metallic oil feed. Thus significant numbers ofliquid hydrocarbon molecules can diffuse to active catalytic sites bothin the matrix and in sieve particles on the surface of the matrix. Ingeneral it is preferred to employ catalysts with matrices wherein thefeeder pores have diameters in the range of about 400 to about 6000angstrom units, and preferably about 1000 to about 6000 angstrom units.

It is considered an advantage that the process of the present inventioncan be conducted in the substantial absence of tin and/or antimony or atleast in the presence of a catalyst which is substantially free ofeither or both of these metals.

The process of the present invention may be operated with the abovedescribed carbo-metallic oil and catalyst as substantially the solematerials charged to the reaction zone. But the charging of additionalmaterials is not excluded. The charging of recycled oil to the reactionzone has already been mentioned. As described in greater detail below,still other materials fulfilling a variety of functions may also becharged. In such case, the carbo-metallic oil and catalyst usuallyrepresent the major proportion by weight of the total of all materialscharged to the reaction zone. Certain of the additional materials whichmay be used perform functions which offer significant advantages overthe process as performed with only the carbo-metallic oil and catalyst.Among these functions are: controlling the effects of heavy metals andother catalyst contaminants; enhancing catalyst activity; absorbingexcess heat in the catalyst as received from the regenerator; disposalof pollutants or conversion thereof to a form or forms in which they maybe more readily separated from products and/or disposed of; controllingcatalyst temperature; diluting the carbo-metallic oil vapors to reducetheir partial pressure and increase the yield of desired products;adjusting feed/catalyst contact time; donation of hydrogen to a hydrogendeficient carbo-metallic oil feedstock, for example as disclosed in aU.S. application Ser. No. 256,791 entitled "Use of Naphtha inCarbo-Metallic Oil Conversion" and filed in the name of George D. Myerson Mar. 23, 1981, assisting in the dispersion of the feed; and possiblyalso distillation of products. Certain of the metals in the heavy metalsaccumulation on the catalyst are more active in promoting undesiredreactions when they are in the form of elemental metal, than they arewhen in the oxidized form produced by contact with oxygen in thecatalyst regenerator. However, the time of contact between catalyst andvapors of feed and product in past conventional catalytic cracking wassufficient so that hydrogen released in the cracking reaction was ableto reconvert a significant portion of the less harmful oxides back tothe more harmful elemental heavy metals. One can take advantage of thissituation through the introduction of additional materials which are ingaseous (including vaporous) form in the reaction zone in admixture withthe catalyst and vapors of feed and products. The increased volume ofmaterial in the reaction zone resulting from the presence of suchadditional materials tends to increase the velocity of flow through thereaction zone with a corresponding decrease in the residence time of thecatalyst and oxidized heavy metals borne thereby. Because of thisreduced residence time, there is less opportunity for reduction of theoxidized heavy metals to elemental form and therefore less of theharmful elemental metals are available for contacting the feed andproducts.

Added materials may be introduced into the process in any suitablefashion, some examples of which follow. For instance, they may beadmixed with the carbo-metallic oil feedstock prior to contact of thelatter with the catalyst. Alternatively, the added materials may, ifdesired, be admixed with the catalyst prior to contact of the latterwith the feedstock. Separate portions of the added materials may beseparately admixed with both catalyst and carbo-metallic oil. Moreover,the feedstock, catalyst and additional materials may, if desired, bebrought together substantially simultaneously. A portion of the addedmaterials may be mixed with catalyst and/or carbo-metallic oil in any ofthe above described ways, while additional portions are subsequentlybrought into admixture. For example, a portion of the added materialsmay be added to the carbo-metallic oil and/or to the catalyst beforethey reach the reaction zone, while another portion of the addedmaterials is introduced directly into the reaction zone. The addedmaterials may be introduced at a plurality of spaced locations in thereaction zone or along the length thereof, if elongated.

The amount of additional materials which may be present in the feed,catalyst or reaction zone for carrying out the above functions, andothers, may be varied as desired; but said amount will preferably besufficient to substantially heat balance the process. These materialsmay for example be introduced into the reaction zone in a weight ratiorelative to feed of up to about 0.4, preferably in the range of about0.02 to about 0.4, more preferably about 0.03 to about 0.3 and mostpreferably about 0.05 to about 0.25.

For example, many or all of the above desirable functions may beattained by introducing H₂ O to the reaction zone in the form of steamor of liquid water or a combination thereof in a weight ratio relativeto feed in the range of about 0.04 or more, or more preferably about0.05 to about 0.1 or more. Without wishing to be bound by any theory, itappears that the use of H₂ O tends to inhibit reduction ofcatalyst-borne oxides, sulfites and sulfides to the free metallic formwhich is believed to promote condensation-dehydrogenation withconsequent promotion of coke and hydrogen yield and accompanying loss ofproduct. Moreover, H₂ O may also, to some extent, reduce deposition ofmetals onto the catalyst surface. There may also be some tendency todesorb nitrogen-containing and other heavy contaminant-containingmolecules from the surface of the catalyst particles, or at least sometendency to inhibit their absorption by the catalyst. It is alsobelieved that added H₂ O tends to increase the acidity of the catalystby Bronsted acid formation which in turn enhances the activity of thecatalyst. Assuming the H₂ O as supplied is cooler than the regeneratedcatalyst and/or the temperature of the reaction zone, the sensible heatinvolved in raising the temperature of the H₂ O upon contacting thecatalyst in the reaction zone or elsewhere can absorb excess heat fromthe catalyst. Where the H₂ O is or includes recycled water that containsfor example about 500 to about 5000 ppm of H₂ S dissolved therein, anumber of additional advantages may accrue. The ecologicallyunattractive H₂ S need not be vented to the atmosphere, the recycledwater does not require further treatment to remove H₂ S and the H₂ S maybe of assistance in reducing coking of the catalyst by passivation ofthe heavy metals, i.e. by conversion thereof to the sulfide form whichhas a lesser tendency than the free metals to enhance coke and hydrogenproduction. In the reaction zone, the presence of H₂ O can dilute thecarbo-metallic oil vapors, thus reducing their partial pressure andtending to increase the yield of the desired products. It has beenreported that H₂ O is useful in combination with other materials ingenerating hydrogen during cracking; thus it may be able to act as ahydrogen donor for hydrogen deficient carbo-metallic oil feedstocks. TheH₂ O may also serve certain purely mechanical functions such as:assisting in the atomizing or dispersion of the feed; competing withhigh molecular weight molecules for adsorption on the surface of thecatalyst, thus interrupting coke formation; steam distillation ofvaporizable product from unvaporized feed material; and disengagement ofproduct from catalyst upon conclusion of the cracking reaction. It isparticularly preferred to bring together H₂ O, catalyst andcarbo-metallic oil substantially simultaneously. For example, one mayadmix H₂ O and feedstock in an atomizing nozzle and immediately directthe resultant spray into contact with the catalyst at the downstream endof the reaction zone.

The addition of steam to the reaction zone is frequently mentioned inthe literature of fluid catalytic cracking. Addition of liquid water tothe feed is discussed relatively infrequently, compared to theintroduction of steam directly into the reaction zone. However, inaccordance with the present invention it is particularly preferred thatliquid water be brought into intimate admixture with the carbo-metallicoil in a weight ratio of about 0.04 to about 0.25 at or prior to thetime of introduction of the oil into the reaction zone, whereby thewater (e.g., in the form of liquid water or in the form of steamproduced by vaporization of liquid water in contact with the oil) entersthe reaction zone as part of the flow of feedstock which enters suchzone. Although not wishing to be bound by any theory, it is believedthat the foregoing is advantageous in promoting dispersion of thefeedstock. Also, the heat of vaporization of the water, which heat isabsorbed from the catalyst, from the feedstock, or from both, causes thewater to be a more efficient heat sink than steam alone. Preferably theweight ratio of liquid water to feed is about 0.04 to about 0.2 morepreferably about 0.05 to about 0.15.

Of course, the liquid water may be introduced into the process in theabove described manner or in other ways, and in either event theintroduction of liquid water may be accompanied by the introduction ofadditional amounts of water as steam into the same or different portionsof the reaction zone or into the catalyst and/or feedstock. For example,the amount of additional steam may be in a weight ratio relative to feedin the range of about 0.01 to about 0.25, with the weight ratio of totalH₂ O (as steam and liquid water) to feedstock being about 0.3 or less.The charging weight ratio of liquid water relative to steam in suchcombined use of liquid water and steam may for example range from about15 which is presently preferred, to about 0.2. Such ratio may bemaintained at a predetermined level within such range or varied asnecessary or desired to adjust or maintain heat balance.

Other materials may be added to the reaction zone to perform one or moreof the above described functions. For example, thedehydrogenation-condensation activity of heavy metals may be inhibitedby introducing hydrogen sulfide gas into the reaction zone. Hydrogen maybe made available for hydrogen deficient carbo-metallic oil feedstocksby introducing into the reaction zone either a conventional hydrogendonor diluent such as a heavy naphtha or relatively low molecular weightcarbon-hydrogen fragment contributors, including for example: lightparaffins; low molecular weight alcohols and other compounds whichpermit or favor intermolecular hydrogen transfer; and compounds thatchemically combine to generate hydrogen in the reaction zone such as byreaction of carbon monoxide with water, or with alcohols, or witholefins, or with other materials or mixtures of the foregoing.

All of the above mentioned additional materials (including water), aloneor in conjunction with each other or in conjunction with othermaterials, such as nitrogen or other inert gases, light hydrocarbons,and others, may perform any of the above-described functions for whichthey are suitable, including without limitation, acting as diluents toreduce feed partial pressure and/or as heat sinks to absorb excess heatpresent in the catalyst as received from the regeneration step. Theforegoing is a discussion of some of the functions which can beperformed by materials other than catalyst and carbo-metallic oilfeedstock introduced into the reaction zone, and it should be understoodthat other materials may be added or other functions performed withoutdeparting from the spirit of the invention.

The invention may be practiced in a wide variety of apparatus. However,the preferred apparatus includes means for rapidly vaporizing as muchfeed as possible and efficiently admixing feed and catalyst (althoughnot necessarily in that order), for causing the resultant mixture toflow as a dilute suspension in a progressive flow mode, and forseparating the catalyst from cracked products and any uncracked or onlypartially cracked feed at the end of a predetermined residence time ortimes, it being preferred that all or at least a substantial portion ofthe product should be abruptly separated from at least a portion of thecatalyst.

For example, the apparatus may include, along its elongated reactionchamber, one or more points for introduction of carbo-metallic feed, oneor more points for introduction of catalyst, one or more points forintroduction of additional material, one or more points for withdrawalof products and one or more points for withdrawal of catalyst. The meansfor introducing feed, catalyst and other material may range from openpipes to sophisticated jets or spray nozzles, it being preferred to usemeans capable of breaking up the liquid feed into fine droplets.Preferably, the catalyst, liquid water (when used) and fresh feed arebrought together in an apparatus similar to that disclosed in U.S.patent application Ser. No. 969,601 of George D. Myers et al, filed Dec.14, 1978, the entire disclosure of which is hereby incorporated hereinby reference. According to a particularly preferred embodiment based ona suggestion which is understood to have emanated from Mr. Steven M.Kovach, the liquid water and carbo-metallic oil, prior to theirintroduction into the riser, are caused to pass through a propeller,apertured disc, or any appropriate high shear agitating means forforming a "homogenized mixture" containing finely divided droplets ofoil and/or water with oil and/or water present as a continuous phase.

It is preferred that the reaction chamber, or at least the major portionthereof, be more nearly vertical than horizontal and have a length todiameter ratio of at least about 10, more preferably about 20 or 25 ormore. Use of a vertical riser type reactor is preferred. If tubular, thereactor can be of uniform diameter throughout or may be provided with acontinuous or step-wise increase in diameter along the reaction path tomaintain or vary the velocity along the flow path.

In general, the charging means (for catalyst and feed) and the reactorconfiguration are such as to provide a relatively high velocity of flowand dilute suspension of catalyst. For example, the vapor or catalystvelocity in the riser will be usually at least about 25 and moretypically at least about 35 feet per second. This velocity may range upto about 55 or about 75 feet or about 100 feet per second or higher. Thevapor velocity at the top of the reactor may be higher than that at thebottom and may for example be about 80 feet per second at the top andabout 40 feet per second at the bottom. The velocity capabilities of thereactor will in general be sufficient to prevent substantial build-up ofcatalyst bed in the bottom or other portions of the riser, whereby thecatalyst loading in the riser can be maintained below about 4 or 5pounds, as for example about 0.5 pounds, and below about 2 pounds, asfor example 0.8 pound, per cubic foot, respectively, at the upstream(e.g. bottom) and downstream (e.g. top) ends of the riser.

The progressive flow mode involves, for example, flowing of catalyst,feed and products as a stream in a positively controlled and maintaineddirection established by the elongated nature of the reaction zone. Thisis not to suggest however that there must be strictly linear flow. As iswell known, turbulent flow and "slippage" of catalyst may occur to someextent especially in certain ranges of vapor velocity and some catalystloadings, although it has been reported adviseable to employsufficiently low catalyst loadings to restrict slippage and back-mixing.

Most preferably the reactor is one which abruptly separates asubstantial portion or all of the vaporized cracked products from thecatalyst at one or more points along the riser, and preferably separatessubstantially all of the vaporized cracked products from the catalyst atthe downstream end of the riser. A preferred type of reactor embodiesballistic separation of catalyst and products; that is, catalyst isprojected in a direction established by the riser tube, and is caused tocontinue its motion in the general direction so established, while theproducts, having lesser momentum, are caused to make an abrupt change ofdirection, resulting in an abrupt, substantially instantaneousseparation of product from catalyst. In a preferred embodiment referredto as a vented riser, the riser tube is provided with a substantiallyunobstructed discharge opening at its downstream end for discharge ofcatalyst. An exit port in the side of the tube adjacent the downstreamend receives the products. The discharge opening communicates with acatalyst flow path which extends to the usual stripper and regenerator,while the exit port communicates with a product flow path which issubstantially or entirely separated from the catalyst flow path andleads to separation means for separating the products from therelatively small portion of catalyst, if any, which manages to gainentry to the product exit port. Examples of a ballistic separationapparatus and technique as above described, are found in U.S. Pat. Nos.4,066,533 and 4,070,159 to Myers et al, the disclosures of which patentsare hereby incorporated herein by reference in their entireties.According to a particularly preferred embodiment, based on a suggestionunderstood to have emanated from Paul W. Walters, Roger M. Benslay andDwight F. Barger, the ballistic separation step includes at least apartial reversal of direction by the product vapors upon discharge fromthe riser tube; that is, the product vapors make a turn or change ofdirection which exceeds 90° at the riser tube outlet. This may beaccomplished for example by providing a cup-like member surrounding theriser tube at its upper end, the ratio of cross-sectional area of thecup-like member relative to the cross-sectional area of the riser tubeoutlet being low i.e. less than 1 and preferably less than about 0.6.Preferably the lip of the cup is slightly downstream of, or above thedownstream end or top of the riser tube, and the cup is preferablyconcentric with the riser tube. By means of a product vapor linecommunicating with the interior of the cup but not the interior of theriser tube, having its inlet positioned within the cup interior in adirection upstream of the riser tube outlet, product vapors emanatingfrom the riser tube and entering the cup by reversal of direction aretransported away from the cup to catalyst and product separationequipment. Such an arrangement can produce a high degree of completionof the separation of catalyst from product vapors at the riser tubeoutlet, so that the required amount of auxiliary catalyst separationequipment such as cyclones is greatly reduced, with consequent largesavings in capital investment and operating cost.

Preferred conditions for operation of the process are described below.Among these are feed, catalyst and reaction temperatures, reaction andfeed pressures, residence time and levels of conversion, coke productionand coke laydown on catalyst.

In conventional FCC operations with VGO, the feedstock is customarilypreheated, often to temperatures significantly higher than are requiredto make the feed sufficiently fluid for pumping and for introductioninto the reactor. For example, preheat temperatures as high as about700° or 800° F. have been reported. But in our process as presentlypracticed it is preferred to restrict preheating of the feed, so thatthe feed is capable of absorbing a larger amount of heat from thecatalyst while the catalyst raises the feed to conversion temperature,at the same time minimizing utilization of external fuels to heat thefeedstock. Thus, where the nature of the feedstock permits, it may befed at ambient temperature. Heavier stocks may be fed at preheattemperatures of up to about 600° F., typically about 200° F. to about500° F., but higher preheat temperatures are not necessarily excluded.

The catalyst fed to the reactor may vary widely in temperature, forexample from about 1100° to about 1600° F., more preferably about 1200°to about 1500° F. and most preferably about 1300° to about 1400° F.,with about 1325° to about 1375° being considered optimum at present.

As indicated previously, the conversion of the carbo-metallic oil tolower molecular weight products may be conducted at a temperature ofabout 900° to about 1400° F., measured at the reaction chamber outlet.The reaction temperature as measured at said outlet is more preferablymaintained in the range of about 965° to about 1300° F., still morepreferably about 975° to about 1200° F., and most preferably about 980°to about 1150° F. Depending upon the temperature selected and theproperties of the feed, all of the feed may or may not vaporize in theriser.

Although the pressure in the reactor may, as indicated above, range fromabout 10 to about 50 psia, preferred and more preferred pressure rangesare about 15 to about 35 and about 20 to about 35. In general, thepartial (or total) pressure of the feed may be in the range of about 3to about 30, more preferably about 7 to about 25 and most preferablyabout 10 to about 17 psia. The feed partial pressure may be controlledor suppressed by the introduction of gaseous (including vaporous)materials into the reactor, such as for instance the steam, water andother additional materials described above. The process has for examplebeen operated with the ratio of feed partial pressure relative to totalpressure in the riser in the range of about 0.2 to about 0.8, moretypically about 0.3 to about 0.7 and still more typically about 0.4 toabout 0.6, with the ratio of added gaseous material (which may includerecycled gases and/or steam resulting from introduction of H₂ O to theriser in the form of steam and/or liquid water) relative to totalpressure in the riser correspondingly ranging from about 0.8 to about0.2, more typically about 0.7 to about 0.3 and still more typicallyabout 0.6 to about 0.4. In the illustrative operations just described,the ratio of the partial pressure of the added gaseous material relativeto the partial pressure of the feed has been in the range of about 0.25to about 4.0, more typically about 0.4 to about 2.3 and still moretypically about 0.7 to about 1.7.

Although the residence time of feed and product vapors in the riser maybe in the range of about 0.5 to about 10 seconds, as described above,preferred and more preferred values are about 0.5 to about 6 and about 1to about 4 seconds, with about 1.5 to about 3.0 seconds currently beingconsidered about optimum. For example, the process has been operatedwith a riser vapor residence time of about 2.5 seconds or less byintroduction of copious amounts of gaseous materials into the riser,such amounts being sufficient to provide for example a partial pressureratio of added gaseous materials relative to hydrocarbon feed of about0.8 or more. By way of further illustration, the process has beenoperated with said residence time being about two seconds or less, withthe aforesaid ratio being in the range of about 1 to about 2. Thecombination of low feed partial pressure, very low residence time andballistic separation of products from catalyst is considered especiallybeneficial for the conversion of carbo-metallic oils. Additionalbenefits may be obtained in the foregoing combination when there is asubstantial partial pressure of added gaseous material, especially H₂ Oas described above.

Depending upon whether there is slippage between the catalyst andhydrocarbon vapors in the riser, the catalyst riser residence time mayor may not be the same as that of the vapors. Thus, the ratio of averagecatalyst reactor residence time versus vapor reactor residence time,i.e. slippage, may be in the range of about 1 to about 5, morepreferably about 1 to about 4 and most preferably about 1 to about 3,with about 1 to about 2 currently being considered optimum.

In practice, there will usually be a small amount of slippage, e.g., atleast about 1.05 or 1.2. In an operating unit there may for example be aslippage of about 1.1 at the bottom of the riser and about 1.05 at thetop.

In certain types of known FCC units, there is a riser which dischargescatalyst and product vapors together into an enlarged chamber, usuallyconsidered to be part of the reactor, in which the catalyst isdisengaged from product and collected. Continued contact of catalyst,uncracked feed (if any) and cracked products in such enlarged chamberresults in an overall catalyst feed contact time appreciably exceedingthe riser tube residence times of the vapors and catalysts. Whenpracticing the process of the present invention with ballisticseparation of catalyst and vapors at the downstream (e.g. upper)extremity of the riser, such as is taught in the above mentioned Myerset al patents, the riser residence time and the catalyst contact timeare substantially the same for a major portion of the feed and productvapors. It is considered advantageous if the vapor riser residence timeand vapor catalyst contact time are substantially the same for at leastabout 80%, more preferably at least about 90% and most preferably atleast about 95% by volume of the total feed and product vapors passingthrough the riser. By denying such vapors continued contact withcatalyst in a catalyst disengagement and collection chamber one mayavoid a tendency toward re-cracking and diminished selectivity.

In general, the combination of catalyst to oil ratio, temperatures,pressures and residence times should be such as to effect a substantialconversion of the carbo-metallic oil feedstock. It is an advantage ofthe process that very high levels of conversion can be attained in asingle pass; for example the conversion may be in excess of 50% and mayrange to about 90% or higher. Preferably, the aforementioned conditionsare maintained at levels sufficient to maintain conversion levels in therange of about 60 to about 90% and more preferably about 70 to about85%. The foregoing conversion levels are calculated by subtracting from100% the percentage obtained by dividing the liquid volume of fresh feedinto 100 times the volume of liquid product boiling at and above 430° F.(tbp, standard atmospheric pressure).

These substantial levels of conversion may and usually do result inrelatively large yields of coke, such as for example about 4 to about14% by weight based on fresh feed, more commonly about 6 to about 13%and most frequently about 10 to about 13%. The coke yield can more orless quantitatively deposit upon the catalyst. At contemplated catalystto oil ratios, the resultant coke laydown may be in excess of about 0.3,more commonly in excess of about 0.5 and very frequently in excess ofabout 1% of catalyst by weight, based on the weight of moisture freeregenerated catalyst. Such coke laydown may range as high as about 2%,or about 3%, or even higher.

In common with conventional FCC operations on VGO, the present processincludes stripping of spent catalyst after disengagement of the catalystfrom product vapors. Persons skilled in the art are acquainted withappropriate stripping agents and conditions for stripping spentcatalyst, but in some cases the present process may require somewhatmore severe conditions than are commonly employed. This may result, forexample, from the use of a carbo-metallic oil having constituents whichdo not volatilize under the conditions prevailing in the reactor, whichconstituents deposit themselves at least in part on the catalyst. Suchadsorbed, unvaporized material can be troublesome from at least twostandpoints. First, if the gases (including vapors) used to strip thecatalyst can gain admission to a catalyst disengagement or collectionchamber connected to the downstream end of the riser, and if there is anaccumulation of catalyst in such chamber, vaporization of theseunvaporized hydrocarbons in the stripper can be followed by adsorptionon the bed of catalyst in the chamber. More particularly, as thecatalyst in the stripper is stripped of adsorbed feed material, theresultant feed material vapors pass through the bed of catalystaccumulated in the catalyst collection and/or disengagement chamber andmay deposit coke and/or condensed material on the catalyst in said bed.As the catalyst bearing such deposits moves from the bed and into thestripper and from thence to the regenerator, the condensed products cancreate a demand for more stripping capacity, while the coke can tend toincrease regeneration temperatures and/or demand greater regenerationcapacity. For the foregoing reasons, it is preferred to prevent orrestrict contact between stripping vapors and catalyst accumulations inthe catalyst disengagement or collection chamber. This may be done forexample by preventing such accumulations from forming, e.g. with theexception of a quantity of catalyst which essentially drops out ofcirculation and may remain at the bottom of the disengagement and/orcollection chamber, the catalyst that is in circulation may be removedfrom said chamber promptly upon settling to the bottom of the chamber.Also, to minimize regeneration temperatures and demand for regenerationcapacity, it may be desirable to employ conditions of time, temperatureand atmosphere in the stripper which are sufficient to reducepotentially volatile hydrocarbon material borne by the stripped catalystto about 10% or less by weight of the total carbon loading on thecatalyst. Such stripping may for example include reheating of thecatalyst, extensive stripping with steam, the use of gases having atemperature considered higher than normal for FCC/VGO operations, suchas for instance flue gas from the regenerator, as well as other refinerystream gases such as hydro-treater off-gas (H₂ S containing), hydrogenand others. For example, the stripper may be operated at a temperatureof about 350° F. using steam at a pressure of about 150 psig and aweight ratio of steam to catalyst of about 0.002 to about 0.003. On theother hand, the stripper may be operated at a temperature of about 1025°F. or higher.

Substantial conversion of carbo-metallic oils to lighter products inaccordance with the invention tends to produce sufficiently large cokeyields and coke laydown on catalyst to require some care in catalystregeneration. In order to maintain adequate activity in zeolite andnon-zeolite catalysts, it is desirable to regenerate the catalyst underconditions of time, temperature and atmosphere sufficient to reduce thepercent by weight of carbon remaining on the catalyst to about 0.25% orless, whether the catalyst bears a large heavy metals accumulation ornot. Preferably this weight percentage is about 0.1% or less and morepreferably about 0.05% or less, especially with zeolite catalysts. Theamounts of coke which must therefore be burned off of the catalysts whenprocessing carbo-metallic oils are usually substantially greater thanwould be the case when cracking VGO. The term coke when used to describethe present invention, should be understood to include any residualunvaporized feed or cracking product, if any such material is present onthe catalyst after stripping.

Regeneration of catalyst, burning away of coke deposited on the catalystduring the conversion of the feed, may be performed at any suitabletemperature in the range of about 1100° to about 1600° F., measured atthe regenerator catalyst outlet. This temperature is preferably in therange of about 1200° to about 1500° F., more preferably about 1275° toabout 1425° F. and optimally about 1325° to about 1375° F. The processhas been operated, for example, with a fluidized regenerator with thetemperature of the catalyst dense phase in the range of about 1300° toabout 1400° F.

In accordance with the invention, regeneration is conducted whilemaintaining the catalyst in one or more fluidized beds in one or morefluidization chambers. Such fluidized bed operations are characterized,for instance, by one or more fluidized dense beds of ebulliatingparticles having a bed density of, for example, about 25 to about 50pounds per cubic foot. Fluidization is maintained by passing gases,including combustion supporting gases, through the bed at a sufficientvelocity to maintain the particles in a fluidized state but at avelocity which is sufficiently small to prevent substantial entrainmentof particles in the gases. For example, the lineal velocity of thefluidizing gases may be in the range of about 0.2 to about 4 feet persecond and preferably about 0.2 to about 3 feet per second. The averagetotal residence time of the particles in the one or more beds issubstantial, ranging for example from about 5 to about 30, morepreferably about 5 to about 20 and still more preferably about 5 toabout 10 minutes. From the foregoing, it may be readily seen that thefluidized bed regeneration of the present invention is readilydistinguishable from the short-contact, low-density entrainment typeregeneration which has been practiced in some FCC operations.

When regenerating catalyst to very low levels of carbon on regeneratedcatalyst, e.g. about 0.1% or less or about 0.05% or less, based on theweight of regenerated catalyst, it is acceptable to burn off at leastabout the last 10% or at least about the last 5% by weight of coke(based on the total weight of coke on the catalyst immediately prior toregeneration) in contact with combustion producing gases containingexcess oxygen. In this connection it is contemplated that some selectedportion of the coke, ranging from all of the coke down to about the last5 or 10% by weight, can be burned with excess oxygen. By excess oxygenis meant an amount in excess of the stoichiometric requirement forburning all of the hydrogen, all of the carbon and all of the othercombustible components, if any, which are present in the above-mentionedselected portion of the coke immediately prior to regeneration. Thegaseous products of combustion conducted in the presence of excessoxygen will normally include an appreciable amount of free oxygen. Suchfree oxygen, unless removed from the by-product gases or converted tosome other form by a means or process other than regeneration, willnormally manifest itself as free oxygen in the flue gas from theregenerator unit. In order to provide sufficient driving force tocomplete the combustion of the coke with excess oxygen, the amount offree oxygen will normally be not merely appreciable but substantial,i.e. there will be a concentration of at least about 2 mole percent offree oxygen in the total regeneration flue gas recovered from theentire, completed regeneration operation. While such technique iseffective in attaining the desired low levels of carbon on regeneratedcatalyst, it has its limitations and difficulties as will becomeapparent from the discussion below.

Heat released by combustion of coke in the regenerator is absorbed bythe catalyst and can be readily retained thereby until the regeneratedcatalyst is brought into contact with fresh feed. When processingcarbo-metallic oils to the relatively high levels of conversion involvedin the present invention, the amount of regenerator heat which istransmitted to fresh feed by way of recycling regenerated catalyst cansubstantially exceed the level of heat input which is appropriate in theriser for heating and vaporizing the feed and other materials, forsupplying the endothermic heat of reaction for cracking, for making upthe heat losses of the unit and so forth. Thus, in accordance with theinvention, the amount of regenerator heat transmitted to fresh feed maybe controlled, or restricted where necessary, within certain approximateranges. The amount of heat so transmitted may for example be in therange of about 500 to about 1200, more particularly about 600 to about900, and more particularly about 650 to about 850 BTUs per pound offresh feed. The aforesaid ranges refer to the combined heat, in BTUs perpound of fresh feed, which is transmitted by the catalyst to the feedand reaction products (between the contacting of feed with catalyst andthe separation of product from catalyst) for supplying the heat ofreaction (e.g. for cracking) and the difference in enthalpy between theproducts and the fresh feed. Not included in the foregoing are the heatmade available in the reactor by the adsorption of coke on the catalyst,nor the heat consumed by heating, vaporizing or reacting recycle streamsand such added materials as water, steam naphtha and other hydrogendonors, flue gases and inert gases, or by radiation and other losses.

One or a combination of techniques may be utilized in this invention forcontrolling or restricting the amount of regeneration heat transmittedvia catalyst to fresh feed. For example, one may add a combustionmodifier to the cracking catalyst in order to reduce the temperature ofcombustion of coke to carbon dioxide and/or carbon monoxide in theregenerator. Moreover, one may remove heat from the catalyst throughheat exchange means, including for example heat exchangers (e.g. steamcoils) built into the regenerator itself, whereby one may extract heatfrom the catalyst during regeneration. Heat exchangers can be built intocatalyst transfer lines, such as for instance the catalyst return linefrom the regenerator to the reactor, whereby heat may be removed fromthe catalyst after it is regenerated. The amount of heat imparted to thecatalyst in the regenerator may be restricted by reducing the amount ofinsulation on the regenerator to permit some heat loss to thesurrounding atmosphere, especially if feeds of exceedingly high cokingpotential are planned for processing; in general, such loss of heat tothe atmosphere is considered economically less desirable than certain ofthe other alternatives set forth herein. One may also inject coolingfluids into portions of the regenerator other than those occupied by thedense bed, for example water and/or steam, whereby the amount of inertgas available in the regenerator for heat absorption and removal isincreased.

Another suitable and preferred technique for controlling or restrictingthe heat transmitted to fresh feed via recycled regenerated catalystinvolves maintaining a specified ratio between the carbon dioxide andcarbon monoxide formed in the regenerator while such gases are in heatexchange contact or relationship with catalyst undergoing regeneration.In general, all or a major portion by weight of the coke present on thecatalyst immediately prior to regeneration is removed in at least onecombustion zone in which the aforesaid ratio is controlled as describedbelow. More particularly, at least the major portion more preferably atleast about 65% and more preferably at least about 80% by weight of thecoke on the catalyst is removed in a combustion zone in which the molarratio of CO₂ to CO is maintained at a level substantially below 5, e.g.about 4 or less. Looking at the CO₂ /CO relationship from the inversestandpoint, it is preferred that the CO/CO₂ molar ratio should be atleast about 0.25 and preferably at least about 0.3 and still morepreferably about 1 or more or even 1.5 or more.

While persons skilled in the art are aware of techniques for inhibitingthe burning of CO to CO₂, it has been suggested that the mole ratio ofCO:CO₂ should be kept less than 0.2 when regenerating catalyst withlarge heavy metal accumulations resulting from the processing ofcarbo-metallic oils, in this connection see for example U.S. Pat. No.4,162,213 to Zrinscak, Sr. et al. In this invention, however, COproduction is increased while catalyst is regenerated to about 0.1%carbon or less, and preferably about 0.05% carbon or less. Moreover,according to a preferred method of carrying out the invention thesub-process of regeneration, as a whole, may be carried out to theabove-mentioned low levels of carbon on regenerated catalyst with adeficiency of oxygen; more specifically, the total oxygen supplied tothe one or more stages of regeneration can be and preferably is lessthan the stoichiometric amount which would be required to burn allhydrogen in the coke to H₂ O and to burn all carbon in the coke to CO₂.If the coke includes other combustibles, the aforementionedstoichiometric amount can be adjusted to include the amount of oxygenrequired to burn them.

Still another particularly preferred technique for controlling orrestricting the regeneration heat imparted to fresh feed via recycledcatalyst involves the diversion of a portion of the heat borne byrecycled catalyst to added materials introduced into the reactor, suchas the water, steam, naphtha, other hydrogen donors, flue gases, inertgases, and other gaseous or vaporizable materials which may beintroduced into the reactor.

The larger the amount of coke which must be burned from a given weightof catalyst, the greater the potential for exposing the catalyst toexcessive temperatures. Many otherwise desirable and useful crackingcatalysts are particularly susceptible to deactivation at hightemperatures, and among these are quite a few of the costly molecularsieve or zeolite types of catalyst. The crystal structures of zeolitesand the pore structures of the catalyst carriers generally are somewhatsusceptible to thermal and/or hydrothermal degradation. The use of suchcatalysts in catalytic conversion processes for carbo-metallic feedscreates a need for regeneration techniques which will not destroy thecatalyst by exposure to highly severe temperatures and steaming. Suchneed can be met by a multi-stage regeneration process which includesconveying spent catalyst into a first regeneration zone and introducingoxidizing gas thereto. The amount of oxidizing gas that enters saidfirst zone and the concentration of oxygen or oxygen bearing gas thereinare sufficient for only partially effecting the desired conversion ofcoke on the catalyst to carbon oxide gases. The partially regeneratedcatalyst is then removed from the first regeneration zone and isconveyed to a second regeneration zone. Oxidizing gas is introduced intothe second regeneration zone to provide a higher concentration of oxygenor oxygen-containing gas than in the first zone, to complete the removalof carbon to the desired level. The regenerated catalyst may then beremoved from the second zone and recycled to the reactor for contactwith fresh feed. An example of such multi-stage regeneration process isdescribed in U.S. patent application Ser. No. 969,602 of George D. Myerset al, filed Dec. 14, 1978, the entire disclosure of which is herebyincorporated herein by reference. Another example may be found in U.S.Pat. No. 2,398,739.

Multi-stage regeneration offers the possibility of combining oxygendeficient regeneration with the control of the CO:CO₂ molar ratio. Thus,about 50% or more, more preferably about 65% to about 95%, and morepreferably about 80% to about 95% by weight of the coke on the catalystimmediately prior to regeneration may be removed in one or more stagesof regeneration in which the molar ratio of CO:CO₂ is controlled in themanner described above. In combination with the foregoing, the last 5%or more, or 10% or more by weight of the coke originally present, up tothe entire amount of coke remaining after the preceding stage or stages,can be removed in a subsequent stage of regeneration in which moreoxygen is present. Such process is susceptible of operation in such amanner that the total flue gas recovered from the entire, completedregeneration operation contains little or no excess oxygen, i.e. on theorder of about 0.2 mole percent or less, or as low as about 0.1 molepercent or less, which is substantially less than the 2 mole percentwhich has been suggested elsewhere. Thus, multi-stage regeneration isparticularly beneficial in that it provides another convenient techniquefor restricting regeneration heat transmitted to fresh feed viaregenerated catalyst and/or reducing the potential for thermaldeactivation, while simultaneously affording an opportunity to reducethe carbon level on regenerated catalyst to those very low percentages(e.g. about 0.1% or less) which particularly enhance catalyst activity.For example, a two-stage regeneration process may be carried out withthe first stage burning about 80% of the coke at a bed temperature ofabout 1300° F. to produce CO and CO₂ in a molar ratio of CO/CO₂ of about1 and the second stage burning about 20% of the coke at a bedtemperature of about 1350° F. to produce substantially all CO₂ mixedwith free oxygen. Use of the gases from the second stage as combustionsupporting gases for the first stage, along with additional airintroduced into the first stage bed, results in an overall CO to CO₂ratio of about 0.6, with a catalyst residence time of about 5 to 15minutes total in the two zones. Moreover, where the regenerationconditions, e.g. temperature or atmosphere, are substantially lesssevere in the second zone than in the first zone (e.g. by at least about10 and preferably at least about 20° F.), that part of the regenerationsequence which involves the most severe conditions is performed whilethere is still an appreciable amount of coke on the catalyst. Suchoperation may provide some protection of the catalyst from the moresevere conditions. A particularly preferred embodiment of the inventionis two-stage fluidized regeneration at a maximum temperature of about1500° F. with a reduced temperature of at least about 10° or 20° F. inthe dense phase of the second stage as compared to the dense phase ofthe first stage, and with reduction of carbon on catalyst to about 0.05%or less or even about 0.025% or less by weight in the second zone. Infact, catalyst can readily be regenerated to carbon levels as low as0.01% by this technique, even though the carbon on catalyst prior toregeneration is as much as about 1%.

In most circumstances, it will be important to insure that no adsorbedoxygen containing gases are carried into the riser by recycled catalyst.Thus, whenever such action is considered necessary, the catalystdischarged from the regenerator may be stripped with appropriatestripping gases to remove oxygen containing gases. Such stripping mayfor instance be conducted at relatively high temperatures, for exampleabout 1350° to about 1370° F., using steam, nitrogen or other inert gasas the stripping gas(es). The use of nitrogen and other inert gases isbeneficial from the standpoint of avoiding a tendency towardhydro-thermal catalyst deactivation which may result from the use ofsteam.

The following comments and discussion relating to metals management,carbon management and heat management may be of assistance in obtainingbest results when operating the invention. Since these remarks are forthe most part directed to what is considered the best mode of operation,it should be apparent that the invention is not limited to theparticular modes of operation discussed below. Moreover, since certainof these comments are necessarily based on theoretical considerations,there is no intention to be bound by any such theory, whether expressedherein or implicit in the operating suggestions set forth hereinafter.

Although discussed separately below, it is readily apparent that metalsmanagement, carbon management and heat management are inter-related andinterdependent subjects both in theory and practice. While coke yieldand coke laydown on catalyst are primarily the result of the relativelylarge quantities of coke precursors found in carbo-metallic oils, theproduction of coke is exacerbated by high metals accumulations, whichcan also significantly affect catalyst performance. Moreover, the degreeof success experienced in metals management and carbon management willhave a direct influence on the extent to which heat management isnecessary. Moreover, some of the steps taken in support of metalsmanagement have proved very helpful in respect to carbon and heatmanagement. As noted previously the presence of a large heavy metalsaccumulation on the catalyst tends to aggravate the problem ofdehydrogenation and aromatic condensation, resulting in increasedproduction of gases and coke for a feedstock of a given Ramsbottomcarbon value. The introduction of substantial quantities of H₂ O intothe reactor, either in the form of steam or liquid water, appears highlybeneficial from the standpoint of keeping the heavy metals in a lessharmful form, i.e. the oxide rather than metallic form. This is ofassistance in maintaining the desired selectivity.

Also, a unit design in which system components and residence times areselected to reduce the ratio of catalyst reactor residence time relativeto catalyst regenerator residence time will tend to reduce the ratio ofthe times during which the catalyst is respectively under reductionconditions and oxidation conditions. This too can assist in maintainingdesired levels of selectivity.

Whether the metals content of the catalyst is being managed successfullymay be observed by monitoring the total hydrogen plus methane producedin the reactor and/or the ratio of hydrogen to methane thus produced. Ingeneral, it is considered that the hydrogen to methane mole ratio shouldbe less than about 1 and preferably about 0.6 or less, with about 0.4 orless being considered about optimum. In actual practice the hydrogen tomethane ratio may range from about 0.5 to about 1.5 and average about0.8 to about 1.

Careful carbon management can improve both selectivity (the ability tomaximize production of valuable products), and heat productivity. Ingeneral, the techniques of metals control described above are also ofassistance in carbon management. The usefulness of water addition inrespect to carbon management has already been spelled out inconsiderable detail in that part of the specification which relates toadded materials for introduction into the reaction zone. In general,those techniques which improve dispersion of the feed in the reactionzone should also prove helpful, these include for instance the use offogging or misting devices to assist in dispersing the feed.

Catalyst to oil ratio is also a factor in heat management. In commonwith prior FCC practice on VGO, the reactor temperature may becontrolled in the practice of the present invention by respectivelyincreasing or decreasing the flow of hot regenerated catalyst to thereactor in response to decreases and increases in reactor temperature,typically the outlet temperature in the case of a riser type reactor.Where the automatic controller for catalyst introduction is set tomaintain an excessive catalyst to oil ratio, one can expectunnecessarily large rates of carbon production and heat release,relative to the weight of fresh feed charged to the reaction zone.

Relatively high reactor temperatures are also beneficial from thestandpoint of carbon management. Such higher temperatures foster morecomplete vaporization of feed and disengagement of product fromcatalyst.

Carbon management can also be facilitated by suitable restriction of thetotal pressure in the reactor and the partial pressure of the feed. Ingeneral, at a given level of conversion, relatively small decreases inthe aforementioned pressures can substantially reduce coke production.This may be due to the fact that restricting total pressure tends toenhance vaporization of high boiling components of the feed, encouragecracking and facilitate disengagement of both unconverted feed andhigher boiling cracked products from the catalyst. It may be ofassistance in this regard to restrict the pressure drop of equipmentdownstream of and in communication with the reactor. But if it isdesired or necessary to operate the system at higher total pressure,such as for instance because of operating limitations (e.g. pressuredrop in downstream equipment) the above described benefits may beobtained by restricting the feed partial pressure. Suitable ranges fortotal reactor pressure and feed partial pressure have been set forthabove, and in general it is desirable to attempt to minimize thepressures within these ranges.

The abrupt separation of catalyst from product vapors and unconvertedfeed (if any) is also of great assistance. It is for this reason thatthe so-called vented riser apparatus and technique disclosed in U.S.Pat. Nos. 4,070,159 and 4,066,533 to George D. Myers et al is thepreferred type of apparatus for conducting this process. For similarreasons, it is beneficial to reduce insofar as possible the elapsed timebetween separation of catalyst from product vapors and the commencementof stripping. The vented riser and prompt stripping tend to reduce theopportunity for coking of unconverted feed and higher boiling crackedproducts adsorbed on the catalyst.

A particularly desirable mode of operation from the standpoint of carbonmanagement is to operate the process in the vented riser using ahydrogen donor if necessary, while maintaining the feed partial pressureand total reactor pressure as low as possible, and incorporatingrelatively large amounts of water, steam and if desired, other diluents,which provide the numerous benefits discussed in greater detail above.Moreover, when liquid water, steam, hydrogen donors, hydrogen and othergaseous or vaporizable materials are fed to the reaction zone, thefeeding of these materials provides an opportunity for exercisingadditional control over catalyst to oil ratio. Thus, for example, thepractice of increasing or decreasing the catalyst to oil ratio for agiven amount of decrease or increase in reactor temperature may bereduced or eliminated by substituting either appropriate reduction orincrease in the charging ratios of the water, steam and other gaseous orvaporizable material, or an appropriate reduction or increase in theratio of water to steam and/or other gaseous materials introduced intothe reaction zone.

Heat management includes measures taken to control the amount of heatreleased in various parts of the process and/or for dealing successfullywith such heat as may be released. Unlike conventional FCC practiceusing VGO, wherein it is usually a problem to generate sufficient heatduring regeneration to heat balance the reactor, the processing ofcarbometallic oils generally produces so much heat as to require carefulmanagement thereof.

Heat management can be facilitated by various techniques associated withthe materials introduced into the reactor. Thus, heat absorption by feedcan be maximized by minimum preheating of feed, it being necessary onlythat the feed temperature be high enough so that it is sufficientlyfluid for successful pumping and dispersion in the reactor. When thecatalyst is maintained in a highly active state with the suppression ofcoking (metals control), so as to achieve higher conversion, theresultant higher conversion and greater selectivity can increase theheat absorption of the reaction. In general, higher reactor temperaturespromote catalyst conversion activity in the face of more refractory andhigher boiling constituents with high coking potentials. While the rateof catalyst deactivation may thus be increased, the higher temperatureof operation tends to offset this loss in activity. Higher temperaturesin the reactor also contribute to enhancement of octane number, thusoffsetting the octane depressant effect of high carbon lay down. Othertechniques for absorbing heat have also been discussed above inconnection with the introduction of water, steam, and other gaseous orvaporizable materials into the reactor.

Another technique employed in heat management is careful stripping. Thesevere stripping techniques referred to above are useful in controllingthe amount of heat released in the regenerator. Each or any combinationof the aforementioned techniques of metals management, carbon managementand heat management may be practiced.

As noted above, the invention can be practised in the abovedescribedmode and in many others. An illustrative, nonlimiting example isdescribed by the accompanying schematic diagrams in the figure and bythe description of this figure which follows.

Referring in detail to the drawing, petroleum feedstock is introducedinto the lower end of riser reactor 2 through inlet line 1, at whichpoint it is mixed with hot regenerated catalyst coming through line 5and stripper 14 from regenerator 9.

The feedstock is catalytically cracked in passing up riser 2 and theproduct vapors are separated from catalyst particles in vessel 3 and areremoved through line 4. The catalyst, contaminated with coke, is passedinto stripper 19 through line 7 and is introduced into bed 23 in upperzone 10 of regenerated 9 through line 36. The rate of flow of catalystinto zone 10 is controlled by valve 8.

Oxidizing gas, such as air, containing about 200 ppm chlorine isintroduced into zone 10 through line 21 and the partially regeneratedcatalyst flows downwardly through conduit 18 into lower regenerationzone 25.

An oxidizing gas, such as air, is introduced into regeneration zone 25through line 11. The oxidizing gas flows through gas distribution plate15 and thus into the bed 16 of catalyst particles. This mixture passesupwardly through the bed 16 of coke-contaminated catalyst particles,fluidizing it as well as reacting with the coke, and passes throughperforated plate 17 into the bed of catalyst particles in zone 10.

The perforations in the plate 17 are large enough so that the upwardlyflowing gas readily passes therethrough into zone 10. Duringregeneration of the catalyst the pressure difference between the upperand lower zones prevents catalyst particles from passing downwardlythrough the plate. Gases within the regenerator comprising combustionproducts, nitrogen, and chlorine, are separated from suspended catalystparticles by a separator (not shown) and then pass out of theregenerator through line 24.

Regenerated catalyst is removed from zone 25 through conduit 26 forreturn to riser 2 through the stripper 14, the rate of removal beingcontrolled by valve 6.

A stripping gas such as steam is introduced into stripper 19 throughline 20 to remove volatiles from the catalyst. The volatiles pass fromthe stripper through line 7 into vessel 3 and then out through line 4.Similarly a stripping gas, such as steam is introduced into stripper 14through line 12 to remove absorbed nitrogen from the regeneratedcatalyst before it is returned to the reactor 2. The stripped gases passthrough line 26 into the regenerator 9.

While this invention may be used with single stage regenerators, or withmultiple stage regenerators having cocurrent instead of countercurrentflow, it is especially useful in a regenerator of the type shown whichis well-suited for producing gases having a high ratio of CO to CO₂.

In a preferred method of carrying out this invention in a countercurrentflow pattern as in the apparatus in the Figure, the amount of oxidizinggas and catalyst are controlled so that the amount of oxidizing gaspassing into zone 25 is greater than that required to convert all thecoke on the catalyst in this zone to carbon dioxide, and the amount ofoxidizing gas passing upwardly from zone 25 into zone 10 together withthe oxidizing gas added to zone 10 from line 21 is insufficient toconvert all the coke in zone 10 to carbon dioxide. Zone 10 thereforewill contain some CO, and Cl₂ add to this zone, or added to a portion ofthe system where the Cl₂ will pass into this zone, will reduce thetendency of CO to oxidize the CO₂. Other portions of the system to whichCl₂ may be added to retard the oxidization of CO to CO₂ include thestripper 14, line 26, bed 16, line 36, and the dilute phase above bed 16and bed 23.

Having thus described this invention, the following Examples are offeredto illustrate the invention in more detail.

EXAMPLE I

A carbo-metallic feed at a temperature of about 450° F. is introduced ata rate of about 2070 pounds per hour into the bottom zone of a ventedriser reactor where it is mixed with steam, water and a zeolite catalystat a temperature of about 1275° F. The catalyst to oil ratio by weightis 11.3 to 1.

The carbo-metallic feed has a heavy metal content of about 5 parts permillion nickel equivalents, and a Conradson carbon content of about 7percent. About 85 percent of the feed boils above 650° F. and about 20%of the feed boils above 1025° F.

The water and steam are injected into the riser at a rate of about 103and 240 pounds per hour respectively. The temperature within the reactoris about 1000° F. and the pressure is about 27 psia. The partialpressure of feed and steam are about 11 psia and 16 psia respectively.

Within the riser about 75 percent of the feed is converted to fractionsboiling at a temperature less than 430° F. and about 53 percent of thefeed is converted to gasoline. During the conversion 11.2 percent of thefeed is converted to coke containing 5.3 percent hydrogen.

The catalyst containing about one percent by weight of coke is removedfrom the reactor and introduced into a stripper where it is contactedwith steam at a temperature of about 1000° F. to remove volatilesadsorbed onto the catalyst. The stripped catalyst is introduced into theupper zone of a two-zone regenerator as shown in the Figure at a rate of23,000 pounds per hour. Each zone contains about 4000 pounds ofcatalyst. Air is introduced into the lower zone at a rate of about 1400pounds per hour. The lower zone produces 85 pounds of CO₂ per hour andno measurable amount of CO, and is at a temperature of about 1340° F.

Air is introduced into the upper zone at a rate of about 1200 pounds perhour together with flue gases from the lower zone. The upper zoneproduces 540 pounds of CO₂ per hour and 112 pounds of CO per hour, andit is at a temperature of about 1330° F.

The regenerator flue gases contain CO₂ and CO in a mol ratio of 3.6. Thecatalyst removed from the lower zone is recycled to the reactor risercontaining about 0.03 percent coke by weight.

EXAMPLE II

The process conditions of Example 1 are followed except that the airadded to the upper zone contains about 130 ppm Cl₂. The flue gascontains CO₂ and CO in a mol ratio of about 2.5. The catalyst removedfrom the lower zone and recycled to the lower riser contains about 0.03percent coke by weight.

The foregoing examples are offered to illustrate this invention and itis obvious that changes may be made in the process without departingfrom the invention.

What is claimed is:
 1. A process for economically convertingcarbo-metallic oils to lighter products comprising:I. providing aconverter feed containing 650° F.+ material, said 650° F.+ materialbeing characterized by a carbon residue on pyrolysis of at least about 1and by containing at least about 4 parts per million of NickelEquivalents of heavy metal(s); II. bringing said converter feed togetherwith particulate cracking catalyst to form a stream comprising asuspension of said catalyst in said feed and causing the resultantstream to flow through a progressive flow type reactor having anelongated reaction chamber which is at least in part vertical orinclined for a predetermined vapor riser residence time in the range ofabout 0.5 to about 10 seconds at a temperature of about 900° to about1400° F. and under a pressure of about 10 to about 50 pounds per squareinch absolute sufficient for causing a conversion per pass in the rangeof about 50% to about 90% while producing coke in amounts in the rangeof about 6 to about 14% by weight based on fresh feed, and laying downcoke on the catalyst in amounts in the range of about 0.3 to about 3% byweight; III. separating spent, coke-laden catalyst from the stream ofhydrocarbons formed by vaporized feed and resultant cracking products;IV. maintaining, in one or more regeneration zones, one or morefluidized catalyst regeneration beds comprising spent catalystundergoing regeneration by combustion of the coke with oxygen on thespent catalyst, and supplying additional spent catalyst to one or moreof such fluidized regeneration bed or beds; V. retaining said catalystparticles in said regeneration zone or zones in contact with a flow ofsaid combustion-supporting gas under conditions of temperature,atmosphere and average total residence time in said zone or zonessufficient for combustion of the coke on the catalyst and for reducingthe level of carbon on the catalyst to about 0.25% by weight or less,while forming gaseous combustion product gases comprising CO and CO₂ ;VI. adding chlorine to the gases within at least one regeneration zonein a concentration high enough to increase the CO/CO₂ ratio; andrecycling the regenerated catalyst to the reactor for contact with freshfeed.
 2. A process according to claim 1 wherein said 650° F.+ materialrepresents at least about 70% by volume of said feed and includes atleast about 10% by volume of material which will not boil below about1000° F.
 3. A process according to claim 1 wherein the carbon residue ofthe feed as a whole corresponds with a Conradson carbon value of atleast about
 2. 4. A process according to claim 1 wherein the carbonresidue of the feed as a whole corresponds with a Conradson carbon valuein the range of about 2 to about
 12. 5. A process according to claim 1wherein the carbon residue of the feed as a whole corresponds with aConradson carbon value of at least about
 6. 6. A process according toclaim 1 wherein the feed as a whole contains at least about 4 parts permillion of Nickel Equivalents of heavy metal present in the form ofelemental metal(s) and/or metal compound(s), of which heavy metal(s) atleast about 2 parts per million is nickel.
 7. A process according toclaim 1 wherein the feed as a whole contains at least about 5.5 partsper million of Nickel Equivalents of heavy metal present in the form ofelemental metal(s) and/or metal compounds(s).
 8. A process according toclaim 1 conducted without prior hydrotreating of the feed.
 9. A processaccording to claim 1 conducted without prior removal of asphaltenes fromthe feed.
 10. A process according to claim 1 conducted without priorremoval of heavy metal(s) from the feed.
 11. A process according toclaim 1 wherein the feed comprises less than about 15% by volume ofrecycled product based on the volume of fresh feed.
 12. A processaccording to claim 1 wherein the catalyst charged to the reactorcomprises an accumulation of heavy metal(s) on said catalyst derivedfrom prior contact under conversion conditions with carbo-metallic oil,said accumulation including about 3000 ppm to about 30,000 ppm of NickelEquivalents of heavy metal(s) and/or metal compound(s) measured inregenerated equilibrium catalyst.
 13. A process according to claim 1wherein the catalyst charged to the reactor is a zeolite molecular sievecatalyst containing at least about 15% by weight of sieve.
 14. A processaccording to claim 1 wherein the catalyst charged to the reactor is azeolite molecular sieve catalyst containing at least about 15% by weightof sieve and comprising an accumulation of heavy metal(s) on saidcatalyst derived from prior contact under conversion conditions withcarbo-metallic oil, said accumulation including about 3000 ppm to about30,000 ppm of Nickel Equivalents of heavy metal(s) by weight, present inthe form of elemental metal(s) and/or metal compound(s), measured inregenerated equilibrium catalyst.
 15. A process according to claim 1wherein make-up catalyst is added to replace catalyst loss or withdrawnfrom the system, said make-up catalyst as introduced having a relativeactivity of at least about 60 percent and any withdrawn catalyst havinga relative activity as withdrawn of at least about 20 percent.
 16. Aprocess according to claim 1 wherein the catalyst has previously beenused to crack a carbo-metallic feed under the conditions recited inclaim
 1. 17. A process according to claim 1 conducted without additionof hydrogen to the reaction zone in which conversion of the feed takesplace.
 18. A process according to claim 1 conducted in the presence, inthe reaction zone, of additional gaseous and/or vaporizable material ina weight ratio, relative to feed, in the range of about 0.02 to about0.4.
 19. A process according to claim 1 wherein the feed is broughttogether with liquid water in a weight ratio relative to feed in therange of about 0.04 to about 0.25 and wherein a stream is formedcontaining a mixture of said feed, said catalyst and steam resultingfrom the vaporization of said liquid water, and is caused to flowthrough said reactor for converting said feed.
 20. A process accordingto claim 19, in which the weight ratio of liquid water to feed is in therange of about 0.05 to about 0.15.
 21. A process according to claim 19in which the water is brought together with the feed at the time of orprior to bringing the feed into contact with the cracking catalyst. 22.A process according to claim 19 in which the water is brought togetherwith the feed prior to bringing the feed into contact with the crackingcatalyst.
 23. A process according to claim 1 wherein the predeterminedriser residence time of the feed and product vapors is about 3 secondsor less.
 24. A process according to claim 1 wherein the temperature insaid reactor is maintained in the range of about 975° F. to about 1200°F.
 25. A process according to claim 1 wherein the temperature in saidreactor is maintained in the range of about 980° F. to about 1150° F.26. A process according to claim 1 wherein the feed partial pressure ismaintained in the range of about 3 to about 30 psia.
 27. A processaccording to claim 1 wherein the feed contains 650° F.+ material whichhas not been hydrotreated and is characterized in part by containing atleast about 5.5 parts per million of Nickel Equivalents of heavymetal(s), present in the form of elemental metal(s) and/or metalcompound(s), said feed being brought together with said crackingcatalyst and with additional gaseous material including steam wherebythe resultant suspension of catalyst and feed also includes gaseousmaterial wherein the ratio of the partial pressure of the added gaseousmaterial relative to the partial pressure of the feed is in the range ofabout 0.25 to about 4.0, and the vapor residence time of feed andproducts in the reactor is in the range of about 0.5 to about 3 seconds.28. A process according to claim 1 wherein all of the feed is cracked inone and the same conversion chamber.
 29. A process according to claim 1wherein the feed is cracked in a substantially single pass mode.
 30. Aprocess according to claim 1 conducted with sufficient severity tomaintain said conversion in the range of about 60 to about 90%.
 31. Aprocess according to claim 1 conducted with sufficient severity tomaintain said conversion in the range of about 70 to about 85%.
 32. Aprocess according to claim 1 wherein at the end of said predeterminedresidence time, the catalyst is projected in a direction established bythe elongated reaction chamber or an extension thereof, while theproducts, having lesser momentum, are caused to make an abrupt change ofdirection relative to the direction in which the catalyst is projected,resulting in an abrupt, substantially instantaneous ballistic separationof products from catalyst.
 33. A process according to claim 1 whereinsaid feed contains 650° F.+ material which has not been hydrotreated andis characterized in part by containing at least about 5.5 parts permillion of Nickel Equivalents of heavy metal(s) present as elementalmetal(s) and/or metal compound(s), said feed being brought together withsaid cracking catalyst and with additional gaseous material includingsteam wherein the resultant suspension of catalyst and feed alsoincludes gaseous material wherein the ratio of the partial pressure ofthe added gaseous material relative to the partial pressure of the feedis in the range of about 0.25 to about 4.0, said vapor residence time offeed and products is in the range of about 0.5 to about 3 seconds andwherein, at the end of said predetermined residence time, the catalystis projected in a direction established by the elongated reactionchamber or an extension thereof, while the products, having lessermomentum, are caused to make an abrupt change of direction relative tothe direction in which the catalyst is projected, resulting in anabrupt, substantially instantaneous ballistic separation of productsfrom catalyst.
 34. A process according to claim 1 wherein theregeneration is conducted in a plurality of regeneration zones.
 35. Aprocess according to claim 1 wherein the regeneration is conducted in aplurality of fluidized catalyst regeneration beds.
 36. A processaccording to claim 1 wherein the regeneration is conducted in aplurality of fluidized catalyst regeneration beds in a plurality ofseparate regeneration zones.
 37. A process according to claim 36 whereinthe spent catalyst and combustion supporting gas are caused to movesequentially through said plurality of zones in directions which are atleast in part countercurrent.
 38. The process of claim 1 wherein thespent catalyst is regenerated in at least two regeneration stages, atleast one of said stages is stoichiometrically deficient in oxygen, andCl₂ is added to said oxygen-deficient stage.
 39. The process of claim 1wherein the spent catalyst is regenerated in at least two regenerationstages, at least one of which is stoichiometrically deficient in oxygenand at least one is oxygen-rich, and Cl₂ is added to said oxygen-richstage.
 40. The process is claim 1 wherein the Conradson carbon contentof the feedstock is greater than about six percent and said Cl₂ is addedto the regenerator in an amount sufficient to provide Cl₂ in aconcentration of from about 100 to about 400 ppm in the gaseous phasewithin said regenerator.
 41. The process of claim 1 wherein theConradson carbon content of the feedstock is in the range of 6 to 12percent and said Cl₂ is added to the generator in an amount sufficientto provide Cl₂ in a concentration of from about 100 to 300 ppm in thegaseous phase within said regenerator.
 42. A process according to claim1 wherein said regeneration is conducted at a temperature in the rangeof about 1100° F. to about 1600° F.
 43. A process according to claim 1wherein said regeneration is conducted at a temperature in the range ofabout 1200° F. to about 1500° F.
 44. A process according to claim 1wherein said regeneration is conducted at a temperature in the range ofabout 1275° F. to about 1425° F.
 45. A process according to claim 1wherein the total residence time of said catalyst particles in saidregeneration zone or zones is in the range of about 5 to about 20minutes.
 46. A process according to claim 1 wherein the total residencetime of said catalyst particles in said regeneration zone or zones is inthe range of about 5 to about 10 minutes.
 47. A process according toclaim 1 wherein CO₂ and CO are formed in a CO₂ to CO mole ratio of nomore than about
 4. 48. A process according to claim 1 wherein the amountof coke removed from said catalyst during regeneration represents about0.5 to about 3% by weight based on the weight of regenerated catalyst.49. A process according to claim 1 wherein the regenerated catalystparticles contain about 0.1% or less by weight of coke.
 50. A processaccording to claim 1 wherein the regenerated catalyst particles containabout 0.05 or less by weight of coke.
 51. A process according to claim 1wherein said catalyst is a zeolite molecular sieve catalyst containingat least about 5% by weight of sieve, the carbon residue of the feed asa while corresponds with a Conradson carbon value of at least about 2,and said 650° F.+ material represents at least about 70% by volume ofsaid feed and includes at least about 10% by volume of material whichwill not boil below about 1000° F.
 52. A process according to claim 51wherein the carbon residue of the feed as a whole corresponds with aConradson carbon value of at least about
 6. 53. A process according toclaim 1 wherein the total residence time of said catalyst particles insaid regeneration zone or zones is in the range of about 5 to about 30minutes.